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PNNL-27186
Conceptual Biorefinery Design and Research Targeted for 2022: Hydrothermal Liquefaction Processing of Wet Waste to Fuels
December 2017
LJ Snowden-Swan RT Hallen
Y Zhu TR Hart
MD Bearden J Liu
TE Seiple KO Albrecht
SB Jones SP Fox
AJ Schmidt GD Maupin
JM Billing DC Elliott
PNNL-27186
Conceptual Biorefinery Design and Research Targeted for 2022: Hydrothermal Liquefaction Processing of Wet Waste to Fuels
LJ Snowden-Swan RT Hallen
Y Zhu TR Hart
MD Bearden J Liu
TE Seiple KO Albrecht
SB Jones SP Fox
AJ Schmidt GD Maupin
JM Billing DC Elliott
December 2017
Prepared for
the U.S. Department of Energy
under Contract DE-AC05-76RL01830
Pacific Northwest National Laboratory
Richland, Washington 99352
PNNL-27186
iii
Summary
The Department of Energy Bioenergy Technologies Office (BETO) invests in research and
development of new pathways for commercially viable conversion of biomass into drop-in ready
transportation fuels, fuel blendstocks and products. The primary emphasis has been on terrestrial and
algae feedstocks, but more recently BETO has begun to explore the potential of wet wastes for biofuel
production, with focus on wastewater residuals, manure, food waste, and fats, oils and grease. A recent
resource analysis estimates that 77 million dry tons per year of these wastes are generated annually, 65%
of which are underutilized for any beneficial purpose.1 Approximately 14 million dry tons of the total
resource is wastewater residuals (sludge and biosolids) generated at the nation’s wastewater treatment
plants (WWTPs).2 Conversion of this resource into transportation fuels could significantly contribute to
the creation of a new domestic bioenergy and bioproduct industry, while providing an economically and
environmentally sustainable alternative for current waste disposal practices.
Hydrothermal liquefaction (HTL) is a process that uses hot, pressurized water in the condensed phase
to convert biomass to a thermally stable oil product, also known as “biocrude”, which can then be thermo-
catalytically upgraded to hydrocarbon fuel blendstocks. HTL is conceptually simple, has a high carbon
efficiency, and can be applied to a wide range of wet feedstocks at similar processing conditions. The
purpose of this report is to document the conceptual design, economics and supporting data for a sludge-
to-fuel pathway via HTL and biocrude upgrading. The configuration includes a HTL plant that is co-
located with a WWTP and a larger scale biocrude upgrading plant for production of hydrocarbon fuel
blendstocks. Experimental data from bench scale testing of a 1:1 mixture of primary:secondary sludges
are used to establish the economic and technical assumptions for the analysis. The design represents a
goal case for the pathway, targeting performance that is anticipated to be achievable by 2022 with further
research and development. The year 2022 is BETO’s target year for verification of hydrocarbon biofuel
pathways.3 As this analysis represents a goal case, assumed values of several design parameters represent
improvements in the technology relative to what has currently been demonstrated in the laboratory.
While HTL is fairly well developed and may therefore be ready for commercialization prior to 2022,
there are specific advancements addressed in this analysis that are necessary to enhance performance
compared to what has been demonstrated to date. In addition, an important aspect to the pathway is the
upgrading of biocrude to fuel blendstock, an area that has received much less attention and requires
significant research to validate the goal case performance parameters.
Summary economics for the sludge HTL plant and the biocrude upgrading plant are presented in
Figures ES.1 and ES.2, respectively. The estimated plant gate minimum fuel selling price for fuel
blendstock from sludge HTL and upgrading is $3.46/gasoline gallon equivalent (gge) (Figure ES.2). This
price is within the tolerance (+$0.49/gge) of BETO’s $3/gge programmatic cost target3 and illustrates that
fuel blendstocks generated from HTL of sludge and centralized biocrude upgrading have the potential to
1 DOE. 2017. Biofuels and Bioproducts from Wet and Gaseous Waste Streams: Challenges and Opportunities.
Bionergy Technologies Office, Energy Efficiency and Renewable Energy, U.S. Department of Energy, Washington,
D.C. 2 Seiple T, AM Coleman, and RL Skaggs. 2017. “Municipal wastewater sludge as a sustainable bioresource in the
United States.” J. Environ. Manage. 197:(2017) 673-680. http://dx.doi.org/10.1016/j.jenvman.2017.04.032 3 DOE. 2016. Biomass Multi-Year Program Plan. Bionergy Technologies Office, Energy Efficiency and
Renewable Energy, U.S. Department of Energy, Washington, D.C.
PNNL-27186
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be competitive with fossil fuels. This analysis illustrates the feasibility of HTL for point-of-generation
conversion of waste feedstock at a scale 1/20th that of the standard lignocellulosic biorefinery scale
typically used in BETO design cases. The relevance of this work reaches beyond wastewater treatment
sludge to lay the groundwork for application to other distributed wet wastes and blends that together
represent a significant resource of underutilized biomass.
Figure ES.1. Summary economics for sludge HTL plant.
Minimum Selling Price $2.35 $/gge biocrude
Minimum Selling Price $2.53 $/gal biocrude
Biocrude 4.1 million gge/yr
0.47 trillion Btu/yr
132 gge/US ton AFDW sludge
15 million Btu/US ton AFDW sludge
Internal Rate of Return (After-Tax) 10%
Equity Percent of Total Investment 40%
Cost Year 2014
Sludge Dewatering $1,400,000 8% Plant Hours per year 7920
HTL Oil Production $13,100,000 72% Feed rate, dry sludge 110 ton/day
HTL Water Recycle Treatment $3,100,000 17% Feed rate, dry ash-free sludge 93.5 ton/day
Balance of Plant $600,000 3% Feed rate, slurry 69 gal/min
Total Installed Capital Cost $18,200,000 100%
$/gge biocrude $/year
Avoided sludge disposal cost 0.00 $0
Building, site development, add'l piping $3,300,000 Natural Gas 0.14 $600,000
Indirect Costs $12,900,000 Chemicals 0.18 $800,000
Working Capital $1,700,000 Electricity 0.11 $500,000
Land (plant located at WWTP) $100,000 Fixed Costs 0.68 $2,800,000
Capital Depreciation 0.27 $1,100,000
Total Capital Investment (TCI) $36,200,000 Average Income Tax 0.16 $700,000
Average Return on Investment 0.81 $3,300,000
Installed Capital per Annual GGE Biocrude $4.5 2.35
TCI per Annual GGE Biocrude $8.9
Net Electricity Purchased (KW) 818
Loan Rate 8.0% Net Electricity Purchased (KWh/gge product) 1.6
Term (years) 10
Capital Charge Factor (computed) 0.141 (Energy in Biocrude) / (Electricity+Natural Gas Input) 4.1
Overall Carbon Yield to Biocrude
On sludge + natural gas 66%
On sludge 72%
PERFORMANCE
Biocrude from Sludge Hydrothermal Liquefaction
CAPITAL COSTS MANUFACTURING COSTS
Primary/Secondary Wastewater Treatment Sludge (25 wt% solids; 15% ash)
PNNL-27186
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Figure ES.2. Summary economics for biocrude upgrading plant.
Biocrude Feedstock Cost: $2.45 $/gge biocrude (includes $0.10/gge transport cost)
Minimum Fuel Selling Price (MFSP) $3.46 $/gge
Diesel Fuel Selling Price $3.71 $/gal
Naphtha Fuel Selling Price $3.42 $/gal
Naphtha Diesel Total
666 1991 2,700 BPSD
9.1 29.6 39 million gge/yr
1.1 3.4 4.5 trillion Btu/yr, LHV basis
0.22 0.73 0.95 gge fuel/gge biocrude
Internal Rate of Return (After-Tax) 10%
Equity Percent of Total Investment 40%
Cost Year 2014
Hydrotreating $33,600,000 45% Plant Hours per year 7920
Hydrocracking $6,600,000 9% Biocrude feed rate 38 mmgal/y
Hydrogen Plant $27,200,000 36%
Steam cycle $1,600,000 2% $/gge fuel blendstock $/year
Balance of Plant $6,500,000 9% Biocrude 2.58 $99,900,000
Total Installed Capital Cost $75,500,000 100% Natural Gas 0.07 $2,800,000
Catalysts & Chemicals 0.01 $500,000
Building, site development, add'l piping $12,400,000 Waste Disposal 0.002 $100,000
Indirect Costs $52,800,000 Electricity and other utilities 0.02 $1,000,000
Working Capital $7,000,000 Fixed Costs 0.24 $9,200,000
Land (included in feedstock cost) $2,700,000 Capital Depreciation 0.002 $4,700,000
Average Income Tax 0.07 $2,800,000
Total Capital Investment (TCI) $150,400,000 Average Return on Investment 0.46 $17,700,000
3.46
Installed Capital per Annual GGE Fuel $2.0
TCI per Annual GGE Fuel $3.9
Net Electricity Purchased (KW) 1,637
Electricity Produced Onsite (KW) 1,812
Electricity Used (KW) 3,449
Loan Rate 8.0% Net Electricity Purchased (KWh/gge product) 0.3
Term (years) 10
Capital Charge Factor (computed) 0.168 Overall Carbon Yield (Naphtha + Diesel)
On biocrude + natural gas 83%
On biocrude 89%
CAPITAL COSTS
Liquid Fuels from Sludge HTL Biocrude Upgrading
MANUFACTURING COSTS
PERFORMANCE
PNNL-27186
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Acknowledgments
The authors gratefully acknowledge the support for this research provided by the U.S. Department of
Energy through the Bioenergy Technologies Office (BETO). We also thank the reviewers for their
valuable feedback and comments on this report. Pacific Northwest National Laboratory is operated for
the U.S. Department of Energy by Battelle under Contract DE-AC06-76RL01830.
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Abbreviations
ACCE Aspen Capital Cost Estimator
AD anaerobic digestion
AFDW ash-free dry weight
BETO Bioenergy Technologies Office
BOD biochemical oxygen demand
BPSD barrels per stream day
Btu British thermal unit
COD chemical oxygen demand
CSTR continuous stirred tank reactor
CWNS Clean Watershed Needs Survey
DOE U.S. Department of Energy
EPA Environmental Protection Agency
FCI fixed capital investment
FOG fats, oils, and grease
gge gasoline gallon equivalent
HHV Higher Heating Value
HTL hydrothermal liquefaction
IRR internal rate of return
LHSV liquid hourly space velocity
LHV lower heating value
MFSP minimum fuel selling price
MGD million gallons per day
PFR plug flow reactor
PNNL Pacific Northwest National Laboratory
POTW publicly owned treatment works
PSA pressure swing adsorption
TAN total acid number
TCI total capital investment
TCLP toxicity characteristic leaching procedure
TDC total direct cost
TEA techno-economic analysis
U overall heat transfer coefficient
USD United States dollar
WAS waste activated sludge
WE&RF Water Environment & Reuse Foundation
WHSV weight hourly space velocity
WWTP waste water treatment plant
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Contents
Summary ...................................................................................................................................................... iii
Acknowledgments ....................................................................................................................................... vii
Abbreviations ............................................................................................................................................... ix
1.0 Introduction .......................................................................................................................................... 1
1.1 Wastewater Treatment Background ............................................................................................. 2
1.2 Overall Sludge to Fuel Process Summary .................................................................................... 4
1.3 Techno-Economic Analysis Approach ......................................................................................... 4
1.3.1 Definition of Nth Plant ...................................................................................................... 5
1.3.2 General Cost Estimation Basis .......................................................................................... 5
2.0 Process Design and Cost Estimation .................................................................................................... 7
2.1 Basis for Sludge Feedstock, Plant Scale, and Sludge Composition ............................................. 7
2.2 Feedstock Preparation .................................................................................................................. 9
2.3 Hydrothermal Liquefaction ........................................................................................................ 10
2.3.1 Sludge HTL Design Basis ............................................................................................... 12
2.3.2 HTL Capital Costs ........................................................................................................... 13
2.4 HTL Aqueous Phase Treatment ................................................................................................. 15
2.4.1 Aqueous Treatment Design Basis ................................................................................... 15
2.4.2 Aqueous Treatment Capital Costs ................................................................................... 19
2.5 Centralized Biocrude Upgrading ................................................................................................ 19
2.5.1 Biocrude Hydroprocessing Design Basis ........................................................................ 20
2.5.2 Biocrude Hydroprocessing Capital Costs ........................................................................ 22
2.5.3 Hydrogen Generation Design and Cost Basis ................................................................. 23
3.0 Process Economics ............................................................................................................................. 25
3.1 Sludge HTL Plant ....................................................................................................................... 25
3.1.1 Total Capital Investment ................................................................................................. 25
3.1.2 Operating Costs ............................................................................................................... 26
3.1.3 Minimum Biocrude Selling Price .................................................................................... 28
3.2 Biocrude Upgrading Plant .......................................................................................................... 29
3.2.1 Total Capital Investment ................................................................................................. 29
3.2.2 Operating Costs ............................................................................................................... 31
3.2.3 Minimum Fuel Selling Price ........................................................................................... 32
4.0 Economic and Technical Sensitivities ................................................................................................ 33
4.1 Sensitivity Analysis .................................................................................................................... 33
4.2 Regional Resource Scale and Fuel Potential Analysis ............................................................... 35
5.0 Conclusions and Recommendations ................................................................................................... 37
6.0 References .......................................................................................................................................... 38
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Appendix A - Heat and Material Balances ............................................................................................... A.1
Appendix B - Equipment Cost Details .......................................................................................................B.1
Appendix C - Compound Selection ...........................................................................................................C.1
Appendix D - Indices ................................................................................................................................ D.1
Appendix E - TCLP Testing Results.......................................................................................................... E.1
Appendix F - Process Efficiency Metrics .................................................................................................. F.1
Figures
Figure 1. Typical steps for conventional biological treatment at a WWTP (adapted from
Tchobanoglous et al. 2014). .................................................................................................................. 3
Figure 2. Simple block diagram for the WWTP/HTL plant and centralized biocrude upgrading
plant. ..................................................................................................................................................... 4
Figure 3. Process flow for slurry feed preparation. .................................................................................... 10
Figure 4. Process flow for sludge HTL. ..................................................................................................... 12
Figure 5. Process flow for HTL aqueous phase treatment. ........................................................................ 18
Figure 6. Process flow for HTL biocrude hydrotreating. ........................................................................... 20
Figure 7. Boiling point distribution (ASTM D2887) for hydrotreated product from sludge
biocrude. ............................................................................................................................................. 22
Figure 8. Process flow for hydrogen production (from Jones et al. 2014). ................................................ 24
Figure 9. Sludge HTL biocrude production cost for goal case. ................................................................. 29
Figure 10. Upgraded fuel production cost for the goal case. ..................................................................... 32
Figure 11. Sensitivity analysis for sludge HTL plant. ............................................................................... 34
Figure 12. Sensitivity analysis for biocrude upgrading plant. ................................................................... 35
Figure 13. Regional fuel production at $3.50/gge from sludge HTL and biocrude upgrading. ................. 36
Tables
Table 1. Nth-plant assumptions. .................................................................................................................. 5
Table 2. Cost factors for direct and indirect costs. ....................................................................................... 6
Table 3. Municipal sludge ash content (% of total solids) ........................................................................... 8
Table 4. Sludge composition and ash content. ............................................................................................. 9
Table 5. Sludge HTL experimental results and model assumptions. ......................................................... 13
Table 6. HTL reactor specifications. .......................................................................................................... 14
Table 7. Sludge HTL capital costs. ............................................................................................................ 15
Table 8. HTL aqueous phase constituent analysis from experimental run ................................................ 16
Table 9. Comparison of HTL effluent water and WWTP headworks component flows. .......................... 17
Table 10. Possible HTL water ammonia removal/recovery options. ......................................................... 19
Table 11. HTL aqueous phase ammonia removal system capital costs ..................................................... 19
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Table 12. Sludge biocrude hydrotreating experimental results and model assumptions. .......................... 21
Table 13. Hydrocracking model assumptions. ........................................................................................... 22
Table 14. Hydrotreater and hydrocracker capital costs. ............................................................................. 23
Table 15. Hydrogen plant model assumptions. .......................................................................................... 24
Table 16. Annual feed and biocrude product rates for the sludge HTL plant. ........................................... 25
Table 17. Total capital investment for sludge HTL plant. ......................................................................... 26
Table 18. HTL plant variable operating costs. ........................................................................................... 27
Table 19. HTL plant fixed operating costs. ............................................................................................... 27
Table 20. Biocrude MFSP cost breakdown. .............................................................................................. 28
Table 21. Annual feed and product flows for centralized biocrude upgrading plant (processing
biocrude from 10 HTL plants). ........................................................................................................... 29
Table 22. Total capital investment for biocrude upgrading plant. ............................................................. 30
Table 23. Biocrude upgrading plant variable operating costs. ................................................................... 31
Table 24. Biocrude upgrading plant fixed operating costs. ....................................................................... 31
Table 25. Upgraded fuel blendstock MFSP cost breakdown. .................................................................... 32
1
1.0 Introduction
For several decades, the Department of Energy Bioenergy Technologies Office (BETO) has
supported research and development for the advancement of bioenergy pathways for terrestrial and algal
feedstocks. More recently, BETO has begun to explore the potential of wet wastes for biofuel production,
and estimates that 77 million dry tons per year of wastewater residuals, manure, food waste, and fats, oils
and grease (FOG) are generated annually, 65% of which are underutilized for any beneficial purpose,
such as for fertilizer, biodiesel or compost (DOE 2017). Conversion of this resource into transportation
fuels could significantly contribute to the nation’s renewable energy goals and provide an economically
and environmentally sustainable alternative for current waste disposal practices.
Approximately 14 million dry tons of wastewater residuals are generated annually at the nation’s
wastewater treatment plants (WWTPs) (Seiple et al. 2017). Costs associated with management and
disposal of these residuals accounts for about 45-65% of the total WWTP’s operating expenses (Nowak
2006; Gray 2010; City of Detroit 2014). Shipping is a significant portion of total sludge management
cost, with some municipalities having to transport their sludge great distances for treatment and disposal
(Peccia and Westerhoff 2015; Hsieh 2013). While cost estimates vary widely, it is evident that reducing
sludge management and disposal costs is a strong industry incentive that introduces the opportunity for
new technology.
Hydrothermal liquefaction (HTL) is a process that uses hot, pressurized water in the condensed phase
to convert wet biomass to an oil product. As such, it is particularly well suited for processing wet waste
feedstocks and eliminates the need for drying that is required for other biomass conversion technologies.
The oil product, also known as “biocrude”, is analogous to petroleum crude in that it contains a mixture of
hydrocarbons with carbon numbers in the gasoline/jet/diesel range. However, it contains higher oxygen
and nitrogen than petroleum and therefore must be hydrotreated to improve compability with petroleum
fuels.
HTL research over the past several decades has included a range of feedstocks from lignocellulosics
to algae to wet wastes. The summary here will focus on liquefaction of sludge feedstock into an oil
product and will not cover the breadth of work done on other feedstocks nor on carbonization of sludge.
Much of the literature on sewage sludge describes testing in lab-scale batch systems. Early work appears
to have originated with Appell et al. (1970), where several organic waste feedstocks were tested in a 500-
ml autoclave system, where the effect of temperature, pressure, and the presence of carbon monoxide (a
reducing agent) and sodium carbonate catalyst on oil yield and composition were examined. A series of
studies in Japan tested various sewage sludges also in the presence of sodium carbonate catalyst
(Yokoyama et al. 1987; Suzuki et al. 1988, 1990; Inoue et al. 1997). More recently, Vardon et al. (2011)
examined properties of biocrude from anaerobically digested sludge, manure and algae. Huang et al.
(2013) and Leng et al. (2015) characterized biocrude from HTL of sludge using ethanol solvent and
studied its emulsification with petroleum diesel, respectively. Malins et al. (2015) examined the effect of
several different catalysts as well as other reaction parameters on biocrude properties. Several recent
studies focused on the migration and ultimate fate of metals during HTL of sludge (Huang et al. 2011;
Huang and Yuan 2016; Leng et al. 2016; Shao et al. 2015; Zhai et al. 2014).
While batch system investigations provide useful information, continuous system testing is necessary
to facilitate engineering scale up and economic analysis of commercial systems. Several studies of
2
continuous HTL focused on sewage sludge can be found in the literature. Appell et al. (1971) conducted
initial trials of a continuous 100-500 g/hr unit developed based on their initial batch testing (Appell et al.
1970) with the use of reducing agents. Solvent extraction was used to separate the oil product from the
reaction mixture. While the focus of this effort was conversion of wastes, including sewage sludge, they
reported results only for cellulose. A study by PNNL for the U.S. Environmental Protection Agency
(EPA) conducted larger scale continuous thermochemical liquefaction as a possible alternative sludge
disposal technology (Molton et al. 1986). In this work, a pilot-scale (30 L/hr) reactor system processing
primary sludge at 20% solids and 5% sodium carbonate concentration was demonstrated and solvent
extraction was used to recover the oil. Itoh et al. (1992) tested a mixture of primary and secondary sludge
with no added catalyst in a 500 kg/day (21 L/hr) continuous system and demonstrated continuous
separation of the oil using high pressure distillation. Villadsen et al. (2012) tested a range of waste
feedstocks including sludge in a 30 L/hr continuous pilot system called CatLiq® (Toor et al. 2012), with
the focus being to develop and apply analysis methodology for identification of specific compounds in
HTL biocrudes. Their process utilizes both homogeneous and heterogeneous catalysts in the reactor, and
a centrifuge for biocrude separation.
As part of a collaborative project between BETO and the Water Environment & Reuse Foundation
(WE&RF), PNNL recently conducted continuous HTL of primary, secondary and anaerobically digested
sludges in a 1.5 L/hr bench system (Marrone 2016). Representatives from WWTPs, the EPA, academia,
and engineering consulting firms participated in the joint project. This work was the basis of a
preliminary techno-economic analysis (TEA) for sludge HTL and subsequent biocrude upgrading
(Snowden-Swan et al. 2016) and helped to guide follow-on experiments with more targeted run
conditions, the results of which provide the basis for the present analysis. Complementing earlier
continuous system studies, PNNL work described in Marrone (2016) and herein provides extensive
characterization of sludge feedstock and products from HTL, introduces potential strategies for reducing
ammonia levels in the HTL water stream recycled to the WWTP, and provides testing and product
characterization data from hydrotreating of sludge-derived biocrude.
The purpose of this report is to document the conceptual design, economics and supporting data for a
sludge HTL plant co-located with a WWTP and a centralized sludge biocrude upgrading plant for
production of hydrocarbon fuel blendstocks. The design and TEA represents a goal case for the pathway,
targeting performance that is anticipated to be achievable by 2022 (BETO’s target year for verification of
biofuel hydrocarbon pathways) with further research and development. As such, values of key design
parameters, such as biocrude yield, sludge feed solids, ash content, and sludge heat transfer rate,
represent improvements in the technology relative to what has currently been demonstrated in the
laboratory. This work builds upon and complements the algae HTL design case (Jones et al. 2014) and
the preliminary sludge HTL TEA (Snowden-Swan et al. 2016) to provide goal case economics for an
example of HTL for onsite, point-of-generation conversion of distributed wet waste. In addition, the
resource and fuel potential analysis illustrates potential regional biocrude collection scenarios that could
enable increased fuel production by taking advantage of the low fuel production cost of the largest scale
WWTPs in the nation.
1.1 Wastewater Treatment Background
Municipal wastewater treatment includes various combinations of process steps, depending on the
constituents needing removal and the levels of removal required (Tchobanoglous et al. 2014). The
3
complete set of steps for taking sewage to clean water for discharge to the environment is commonly
referred to as a “treatment train.” Figure 1 shows a typical treatment train using conventional primary
settling and secondary biological treatment. Preliminary treatment consists of screening of large debris
and settling of grit, both of which are typically landfilled. Primary treatment involves physical settling of
solids from the wastewater while secondary treatment includes biological and/or chemical means to
remove/reduce biodegradable organics, as normally measured by biochemical oxygen demand (BOD) and
total suspended solids (TSS). Primary sludge results from sedimentation of organic solids (e.g., toilet
paper and excrement) from the wastewater. Approximately 50% of the suspended solids and 33% of the
BOD are removed in the primary step (Amuda et al. 2008). Removal of FOG/scum may be carried out in
separate tanks or at the primary sedimentation tanks. The scum is either sent to the plant’s anaerobic
digester (if present) or landfilled. Secondary treatment typically uses aerobic microbes to convert the
remaining organic material into microbial biomass which is then settled and removed in the secondary
clarifiers. The removed solids are commonly referred to as waste activated sludge (WAS). Typical solids
concentrations of primary sludge and secondary sludge are 3% and 0.8%, respectively (Tchobanoglous et
al. 2014). Secondary sludge is often concentrated to 3-4% solids using dissolved air floatation thickeners.
Tertiary treatment with filtration/screening is sometimes used to remove residual suspended solids,
particularly for any water reuse applications. In addition, some plants include nutrient removal in their
secondary or tertiary treatment steps, especially if they are near their permitted nitrogen and/or
phosphorus effluent limits (Tchobanoglous et al. 2014). Nutrient removal may be carried out either
biologically or through chemical methods. More restrictive limits on effluent nutrient levels generally
exist in areas near water bodies that are vulnerable to eutrophication and degraded water quality. The
final step in the liquid treatment train is disinfection, after which the treated water is discharged to the
environment. Various water streams resulting from thickening and dewatering of sludges and digested
biosolids are generally routed back to the plant influent or “headworks” (Tchobanoglous et al. 2014).
Given increasingly strict regulatory limits, the trend is toward more thorough water cleaning to remove a
greater number of constitutents and at higher removal levels (Tchobanoglous et al. 2014). The industry as
a whole is also moving beyond the conventional end-of-pipe treatment role to becoming sustainable
resource recovery facilities of the future, driving toward options that maximize energy and nutrient
recovery while producing clean water.
Screenings and grit to landfill
Primary scum and sludge
Waste activated sludge
Return activated sludge Backwash to headworks
Raw Waste-water
Screening
Grit removal
Primary settling
Aerobic treatment
Secondary settling
Filtration
Effluent toEnvironment
Preliminary TreatmentPrimary
Treatment Conventional Secondary Treatment
TertiaryTreatment(optional) Disinfection
Figure 1. Typical steps for conventional biological treatment at a WWTP (adapted from Tchobanoglous
et al. 2014).
The most common methods that WWTPs use to manage their sludge include stabilization and land
application, landfill disposal, or incineration. Stabilization is commonly achieved through anaerobic
digestion (AD), but can also be accomplished by alkaline addition, drying, or composting. The AD
4
process produces biogas that can be used for onsite heat or power, and “biosolids”, which can be land
applied as a fertilizer alternative. Land application of biosolids is regulated according to CFR 40 Part
503. The type of land to which biosolids may be applied depends on the pathogen and pollutant (metal)
concentrations in the biosolids. Land application of biosolids provides benefit to soils by adding nutrients
and decreasing evaporative water loss, but in some areas faces the challenge of public concern over health
risks (SCAP 2013; Peccia and Westerhoff 2015). HTL could provide a more sustainable alternative than
current sludge management methods, with additional benefits of relatively small footprint, high
operational certainty, high percent solids reduction, and destruction of bioactive components.
1.2 Overall Sludge to Fuel Process Summary
A simple block diagram of the overall HTL plant and biocrude upgrading process scheme is shown
in Figure 2. The HTL facility is assumed to be co-located with the WWTP to avoid the cost of
transporting sludge. It is possible that collection of sludge within a reasonable radius within densely
populated regions may be feasible and should be further investigated to enable larger economies of scale
for the HTL facility. The HTL process produces an oil phase (biocrude), a solids stream containing
mostly ash and some char, and an aqueous stream containing 1-3% carbon. The flow of the aqueous
stream is less than 1% of the headworks flow, however it is highly concentrated in nutrients (relative to
the influent) and therefore may need treatment prior to recycling back to the headworks of the WWTP.
The biocrude is transported by tanker truck to a larger scale centralized upgrading plant for conversion to
fuel naphtha and diesel blendstocks. The use of natural gas is assumed at both the HTL and upgrading
facilities for process heat and hydrogen, respectively.
Figure 2. Simple block diagram for the WWTP/HTL plant and centralized biocrude upgrading plant.
1.3 Techno-Economic Analysis Approach
The approach to developing conversion process techno-economics is similar to that employed in
previous analyses conducted for the Bioenergy Technologies Office (BETO) (Dutta et al. 2015; Jones et
al. 2013, 2014; Tan et al. 2015). Process flow diagrams and models are based on experimental results
from completed and ongoing research, as well as information from commercial vendors for mature and
similar technologies. To assure consistency across all biomass conversion pathways, BETO developed a
set of economic assumptions that are used for all TEAs and are documented in BETO’s Multi-Year
Program Plan (DOE 2016). An important aspect of these assumptions is that they reflect an “nth plant”
design, as described below.
5
1.3.1 Definition of Nth Plant
A standard reference basis common to the conceptual design reports, known as the “nth” plant design,
is used. These assumptions do not account for additional costs that would normally be incurred for a
first-of-a-kind plant, including special financing, equipment redundancies, large contingencies and longer
startup times necessary for the first few plants. For nth plant designs, it is assumed that the costs reflect a
future time when the technology is mature and several plants have already been built and are operating.
The specific assumptions are shown in Table 1. These assumptions are consistent across BETO design
cases, thus allowing a standard basis for comparison of different conversion technologies within the
context of a well-defined hypothetical plant. While WWTPs may use other economic assumptions or
methods of estimating project feasibility when evaluating alternative sludge treatment technologies, this
analysis should provide the key information needed to apply to a specific plant’s circumstances. It is also
worth noting that tax incentives and other credits that may be applicable (e.g., credits under the
Renewable Fuel Standard or cellulosic biofuels bonus depreciation) but are excluded from the analysis to
represent plant economics independent of any government subsidies.
Table 1. Nth-plant assumptions.
Assumption Description Assumed Value
Internal rate of return (IRR) 10%
Plant financing debt/equity 60% / 40% of total capital investment (TCI)
Plant life 30 years
Income tax rate 35%
Interest rate for debt financing 8.0% annually
Term for debt financing 10 years
Working capital cost 5.0% of fixed capital investment (excluding land)
Depreciation schedule 7-years MACRS(a) schedule
Construction period 3 years (8% 1st yr, 60% 2nd yr, 32% 3rd yr)
Plant salvage value No value
Start-up time 6 months
Revenue and costs during start-up Revenue = 50% of normal
Variable costs = 75% of normal
Fixed costs = 100% of normal
On-stream factor 90% (7,884 operating hours per year)
(a) Modified accelerated cost recovery system
1.3.2 General Cost Estimation Basis
All costs in this report are on a 2014 constant dollar basis. This is the current reference year that
BETO uses to facilitate comparison of various conversion technologies (DOE 2016). Indices used to
convert capital and operating costs to 2014 dollars can be found in Appendix D.
Capital costs are estimated from a variety of resources. The heat and material balances generated by
the simulation software (Aspen Plus [AspenTech 2013]) are used to size the major pieces of equipment.
These are input to Aspen Capital Cost Estimator (ACCE) software to determine the installed capital cost.
6
In addition, select data from commercial vendors, either as budgetary estimates or from their published
literature are used when available.
The original cost reflects the year of the cost quote or estimate, and the scale of the equipment. All
capital costs are adjusted to an annualized 2014 basis using the Chemical Engineering (CE) magazine’s
published indices:
Cost in 2014 $ = equipment cost in quote year × (2014 index = 576.1
quote cost year index)
The scale is adjusted to match the appropriate scaling term (heat exchanger area for example) by
using the following expression:
Scaled equipment cost = cost at original scale × (scale up capacity
original capacity)
n
where ‘n’ is the scale factor, typically, 0.6 to 0.7.
Once the equipment is scaled and adjusted to the common cost year, factors are applied to calculate
the total capital investment. Individual installation factors calculated by ACCE are multiplied to
equipment costs, unless installed costs are already available from vendors. The total direct cost is the sum
of all the installed equipment costs, plus the costs for buildings, additional piping and site development.
Indirect costs are estimated as 60% of the total installed costs. Factors for the calculation of these
additional direct and indirect costs are listed in Table 2. The sum of the direct and indirect costs is the
fixed capital investment (FCI). The total capital investment is the fixed capital plus working capital and
land costs.
Table 2. Cost factors for direct and indirect costs.
Direct Costs
Item % of Total Installed Cost (TIC)
Buildings 4.0%
Site development 10.0%
Additional piping 4.5%
Total Direct Costs (TDC) 18.5%
Indirect Costs
Item % of TDC
Prorated expenses 10%
Home office & construction fees 20%
Field expenses 10%
Project contingency 10%
Startup and permits 10%
Total Indirect Costs 60%
Working Capital 5% of FCI
7
Operating costs are estimated by using the results from the Aspen Plus heat and material balances and
applying the assumptions shown in Section 3.0. For the cooling tower, it is assumed that water is
available at 90 °F with a 20 °F allowable temperature rise.
2.0 Process Design and Cost Estimation
This section describes the bases for the design and costing of the sludge HTL plant and the
centralized biocrude upgrading plant. Detailed heat and material balances are presented in Appendix A.
2.1 Basis for Sludge Feedstock, Plant Scale, and Sludge Composition
Sludges from primary and secondary treatment, as well as biosolids from AD have all been shown to
be viable feedstocks for the HTL process (Marrone 2016). A common process configuration includes
both primary and secondary treatment, and therefore, mixed primary/secondary sludge was selected as the
feedstock for the baseline model. While there will be variation from plant to plant in the relative
proportions of primary and secondary sludge generated, a 1:1 ratio was chosen as representative of a
typical plant (Turovskiy and Mathai 2006; ASCE 2000). HTL can also be applied to biosolids from AD,
however direct conversion of sludge obviously results in a greater yield of liquid fuel from the plant’s
influent wastewater carbon.
Existing municipal wastewater treatment plants range from < 0.1 million gallon per day (MGD) to
950 MGD (EPA 2016). The largest 3% of these plants (> 13 MGD) produce approximately 60% of the
total sludge generated in the U.S. (Seiple et al. 2017). As the size of WWTPs is highly variable, the scale
of the HTL and upgrading plant were chosen through an iterative process whereby economics were
evaluated and then sensitivity and regional analyses were performed to support a reasonable base case
scale that is in context of the current WWTP size range and geographical distribution. Through this
process, a base case scale of 110 dry ton/day sludge (including ash) was selected as the approximate
minimum size that is economically feasible (due to economies of scale). This corresponds to a WWTP
that processes about 110 MGD incoming wastewater and serves approximately 1.3 million people (EPA
2016). While this plant scale represents the minimal scale economically feasible for an individual HTL
plant, the regional analysis presented in Section 4.0 explores how collecting biocrude in the most densely
populated areas of the country can make smaller plants viable as part of a larger supply network to a
centralized upgrading facility. It is important to note that the economics presented in this analysis reflect
the cost of conventional on-site construction (“stick-built”) that is typical for chemical plants and
refineries. Alternatively, modularization of HTL systems is possible, which could enable smaller scale
applications through reduced investment costs1. As an example, one company estimated that, through
reduced field costs, modularization resulted in a 15% lower project cost than stick-built construction
1 Weber R, J Holladay, C Jenks, E Panisko, L Snowden-Swan, M Gaalswyck, M Ramirez-Corredores, B Baynes, B
Rittman, M Wright, C Hamstra, L Angenent, SJ Hu, L Harmon, and D Boysen. “Modularized Production of Fuels
and other Value-Added Products from Distributed, Waste or Stranded Feedstocks.” Submitted to WIREs.
8
(Jameson 2007). Further work is needed to investigate and develop methods for estimating the impact of
modularized manufacturing of HTL systems on the plant project economics and MFSP.
The upgrading plant is assumed to receive biocrude from 10 HTL plants within a 100-mile radius and
produces 2,700 barrels per stream day (BPSD) of fuel blendstocks. Further analysis by the resource
assessment team is needed to determine the optimum radius for individual regions in the country. As
illustrated by the regional analysis presented in Section 4.2, this scale and associated biocrude
intermediate availability is reasonable in several of the most densely populated regions of the country.
For perspective, the upgrader scale is 54% of other BETO design cases of ~5,000 BPSD fuel production
and only 5% of the average scale for combined gasoline and diesel production at U.S. refineries of about
51,000 BPSD (EIA 2016).
Sludge treatment and disposal currently represents a significant expense to WWTPs. Therefore, the
solids reduction resulting from the HTL process would conceivably result in significant savings in
avoided disposal costs to the plant. However, the base case model does not include this savings (i.e., a
zero feedstock cost is used) in order to account for potential market adjustments that may result from
implementation of this technology in the future. Sensitivity analysis around this assumption is presented
in Section 4.0.
Wastewater treatment sludge is composed of water, three basic types of organic macromolecules:
proteins, lipids/fats, and carbohydrates, as well as some inorganic material (ash and grit). Sludge
composition will vary from plant to plant depending on the treatment processes used, types of industrial
dischargers in the region (i.e., sewage chemistry), weather, and other factors. Compositional data on
wastewater sludge is sparse in the literature and those that are available are highly variable and often do
not indicate the source of the sludge (e.g., primary, secondary, mixed, or biosolids). While lipid,
carbohydrate and protein content will vary to some degree from plant to plant, ash content is perhaps the
single most important compositional parameter, as it most directly impacts the overall sludge to biocrude
yield for a given plant size. In other words, the ash in the sludge increases equipment size and therefore
capital expense, but does not contribute any fuel product. Table 3 lists ranges of sludge ash content found
in the literature indicating a wide range of 15-40% for mixed primary/secondary.
Table 3. Municipal sludge ash content (% of total solids)
Primary Secondary Mixed Sludge Source(a)
36 14 25 This work
8 16 12(b) Marrone 2016
35 23-33 28 Manara and Zabaniotou 2012
20-30 NR 15-25 Pennsylvania DEP 2001
10-30 25-35 18-33(b) New Mexico Environment Department 2007
15-40 15-40 15-40(b) Tchobanoglous et al. 2014
20-40 15-40 18-40(b) ASCE 2000
20-25 20-25 20-25(b) von Sperling and Gonçalves 2007
20-30 WEF 2013
(a) All sources except this work and Marrone reported as % volatile matter (VM). Ash is calculated as 100-%VM.
(b) Calculated average mix of 1:1 primary, secondary mixture.
9
Ultimate analysis for a 1:1 mixture of primary/secondary sludge provided by the City of Detroit and
Great Lakes Water Authority for HTL testing is given in Table 4. This composition is used for the
modeled sludge feedstock with the exception of ash content, which was decreased from 26% to 15% for
the goal case. A lower ash content was chosen based on two reasons: 1) process knowledge from Detroit
indicates their ash content is likely on the high end of the potential range due to aging infrastructure and
addition of FeCl3 for phosphorus removal in its primary sedimentation step; and 2) the goal case assumes
that future renovations in aging collection systems and outdated processes will lead to reduced sludge ash
content over time. Ideally, enhanced de-gritting and de-ashing technology should also be pursued for
pretreatment of the sludge. The cost of additional processing steps needed for producing a lower ash feed
should be considered in future analyses. Sensitivity analysis around ash content is presented in Section
4.0.
Table 4. Sludge composition and ash content.
50/50 Primary/Secondary
Sludge Mixture Characteristics
Experimental
Data
Model Experimental & Model
Component Wt% dry basis Wt% dry basis Wt% dry, ash free basis
C 41.06 46.83 52.11
H 5.67 6.46 7.19
O 26.06 29.72 33.07
N 4.98 5.68 6.32
S 1.03 1.170 1.30
Ash 26.08 15.0
P 1.86 1.86
HHV Btu/lb(a) 9,936
(a) Calculated by the Boie Equation: HHV (Btu/lb) = (151.2 C + 499.77 H +45.0 S -47.7 O + 27 N)
*100 - 189.0.
2.2 Feedstock Preparation
Dewatering of sludge feedstock is necessary to minimize capital and operating costs for the HTL
process. This section describes the equipment and associated capital necessary to prepare the sludge
slurry feedstock.
Figure 3 shows a simple diagram of the proposed feed preparation steps. Sludges are generated from
the WWTP’s primary and secondary treatment steps at 3% and 0.8% solids content, respectively.
Primary sludge and secondary sludge may be thickened separately before they are mixed as part of
normal WWTP operations. For this analysis, it is assumed that the mixed sludge is intercepted and
mechanically dewatered to the target solids content of 25%. This target is considered the maximum that
is attainable with conventional dewatering technology while still enabling effective pumping through the
HTL system. Secondary sludge is typically more difficult to dewater than primary sludge because of
interstitial water contained in cell mass, therefore mixing the sludge prior to dewatering is likely to
produce the target solids content.
Dewatering of sludges and biosolids is currently carried out to varying degrees at WWTPs to
minimize costs for treatment, transport, and disposition. Common units for dewatering to a high solids
content are the centrifuge, belt filter press, and screw press. Although dewatering equipment are likely to
already be present at the WWTP, capital for investing in a new centrifuge is included in the analysis to
10
account for situations in which the existing dewatering unit is toward the end of its life and/or is not a
large enough capacity to handle raw sludges (e.g., in the case of replacing AD). Chemical conditioning of
the secondary sludge is assumed to be required to maximize dewatering efficiency. Water soluble
polymers, or organic polyelectrolytes, are the most commonly used conditioning agents for mechanical
dewatering and act as flocculants that enhance solid/liquid separation in the sludge (Tchobanoglous et al.
2014). Cationic polyacrylamide is the backbone of most polymer used in the industry (Tchobanoglous et
al. 2014). Polymers provide added benefit over inorganic conditioners by not adding ash to the sludge.
Figure 3. Process flow for slurry feed preparation.
The capital cost for the centrifuge is based on a vendor budgetary estimate. The installed equipment
cost for the centrifuge is $1.4 million (2014 $) for 30,000 gal/hr feed capacity. Individual plants have
differing configurations and therefore slight modifications to capital investment and operating costs are to
be expected. Also, certain configurations of biological treatment, such as the use of membrane
bioreactors in lieu of the conventional activated sludge process, or the use of nutrient removal, make
sludge that is difficult to dewater (Tchobanoglous et al. 2014). Centrifugation of a mixture of primary
sludge/WAS of 2-4% solids content, with added polymer at 5-16 lb/dry ton dry solids, is expected to
achieve 25-35% solids content in the product (Stubbart et al. 2006).
2.3 Hydrothermal Liquefaction
The main process steps in the HTL section of the plant are shown in Figure 4. The base case plant
processes 100 dry tons per day of sludge. The modeled process is largely based on previous conceptual
designs for wood HTL and algae HTL (Knorr et al. 2013; Jones et al. 2014). The 25% solids ground
slurry feed is pumped to 2900-3000 psia and then preheated to 550°F (288°C) in two double-pipe heat
exchangers in series using heat from the reactor liquid product (biocrude/aqueous mixture). High-
pressure pumping of biomass slurries has only been demonstrated at bench and smaller pilot scale,
however, it is generally thought to be more easily accomplished at large scale (Elliott 2011). Berglin et
al. (2012) indicated that pumping a 15% solids wood slurry at HTL pressures is viable with commercial
off-the-shelf equipment, such as a twin-screw feeder/piston pump type of system. Sludge is more easily
pumpable than wood slurry and therefore should be possible at scale and at the higher feed solids
assumed in this analysis. A fired heater with hot oil system is used to bring the pressurized feed up to the
reactor temperature of 656°F (347°C) and provide reactor heat. Off-gas from the reactor along with
purchased natural gas are burned in the fired heater to provide this heat. Scrubbing of the off-gas may be
required, depending on local permit requirements. Approximately 20% of the heat is provided by the
HTL off-gas, which is composed of 92% CO2 and 8% C1-C5 gasses).
Mixed Primary/Secondary Sludge
Dewatered Sludge to HTL
Centrifuge
Polymer
Centrate Return to Headworks
11
The heated pressurized slurry is fed to the HTL reactor where the contents are converted to an organic
biocrude phase, an aqueous phase, solids and and a small amount of gases. Water at subcritical
conditions has a much lower dielectric constant and higher ion product than water at normal conditions
and therefore provides a reaction medium with improved solvent and catalytic properties (Elliott 2011;
Elliott 2015). The HTL chemistry is complicated, however the general reaction pathways can be put into
three basic categories: 1) depolymerization of the biomass components, 2) decomposition of biomass
monomers by cleavage, dehydration, decarboxylization and deamination, and 3) recombination of
reactive fragments (Toor et al. 2012). The biocrude from sludge is similar to biocrude from algae HTL
and comprises a mixture of fatty acids, amides, ketones, hydrocarbons, phenols, alcohols and other
components.
The HTL reactor effluent is fed to a hot filter where solids are removed. The solids stream consists of
60-70% water, ash, char, and low levels of organics from the aqueous phase and biocrude phase. The
biocrude tends to adhere to the solid particles and therefore the amount lost to the solids depends upon the
ash content in the feed. The mass balance from testing indicates that the solid phase from testing of
Detroit sludge (20% feed solids and 26% ash) contained 7-9% of the total produced biocrude. It is
estimated that about half of this biocrude can be recovered through further work on improved separations
methods. The solids are assumed to be disposed of in a landfill. Solids from experimental HTL testing
were analyzed for hazardous metals content with the toxicity characteristic leaching procedure (TCLP)
and determined to be under the federal regulatory limits. While these results indicate that HTL solids are
non-hazardous, the metals content will depend on the incoming sludge composition and therefore TCLP
testing would need to be performed at each plant and probably on a periodic basis to verify that the solids
can be landfilled.. Metals tested include arsenic, barium, cadmium, chromium, copper, lead, mercury,
nickel, selenium, silver, and zinc (see Appendix E). An attractive alternative to landfilling is to recover
value from the HTL solids, either in the form of separate nutrients (e.g., phosphorus) or a combined
fertilizer product. Further work is needed to explore the technical and economical feasibility of this
option.
After solids separation, the remaining biocrude-aqueous-gas mixture is cooled to 140°F (60°C),
reduced to 30 psia, and separated in a three-phase separator. The biocrude product is stored and then
shipped to a centralized upgrading facility where biocrude is assumed to be collected from multiple HTL
plants in the area. The aqueous stream, which contains effluent water, any remaining soluble organics,
ammonia, and metal salts, is routed to the HTL aqueous phase treatment section of the plant (see section
2.4).
12
Heat Exchangers
Filter
Three-Phase Separator
Aqueous to Treatment
Biocrude to Upgrader
FiredPreheaters
Ash+CharHTL Plug Flow
Reactors
Sludge Slurry
Natural Gas
Off Gas
Twin-Screw Feeder / Pump
Figure 4. Process flow for sludge HTL.
2.3.1 Sludge HTL Design Basis
Table 5 shows the experimental testing conditions and product results for the primary/secondary
sludge and the modeled parameters used for the goal case design. The experimental data were collected
from PNNL’s bench scale system similar to previous studies (Elliott et al. 2013; Marrone 2016).
Biocrude yield and feed solids content are important cost drivers that can be improved with further
research. As such, these parameters have been improved for the goal case relative to the experimental
results, with values the experimental team expects can be realized with further research. While there are
potentially thousands of individual compounds present in the biocrude, a simplified list of chemical
compounds is used in the process model. This analysis assumes the same set of compounds modeled in
the previous analysis for algae HTL and biocrude upgrading (Jones et al. 2014), as listed in Appendix C.
This is a reasonable assumption, as biocrude from wastewater sludge has been shown to be similar in
composition to biocrude from algae (Jarvis et al. 2017).
13
Table 5. Sludge HTL experimental results and model assumptions.
Operating Conditions and Results
Experimental Results
(WW06 SS-2)
Aspen Model
Assumption (Goal
Case)
Temperature, °F (°C) 656 (347) 656 (347)
Pressure, psia (MPa) 2979 (20.5) 2979 (20.5)
Feed solids, wt%
Ash included
Ash-free basis
20%
15%
25%
21%
LHSV, vol./h per vol. reactor
Equivalent residence time, minutes
3.6 Hybrid CSTR-PFR(d)
17
6 PFR
10
Product yields(a) (dry, ash free sludge),
wt%
Oil
Aqueous
Gas
Solids
44%
31%
16%
9%
48%
25%
16%
11%
HTL dry oil analysis, wt%
C
H
O
N
S
P
Ash
78.5%
10.7%
4.7%
4.8%
1.2%
0.0
0.06%
78.3%
10.8%
4.8%
4.9%
1.2%
Not modeled(b)
0.0%
HTL dry oil H:C Ratio 1.6 1.6
HTL oil dry HHV, Btu/lb (MJ/kg) 17,000 (39.5)(c) 17,100 (39.7)
HTL oil moisture, wt% 4.4 wt% 4.0 wt%
HTL oil wet density 0.98 0.98
Aqueous phase COD (mg/L) 61,300 61,100
Aqueous phase density (g/ml) 1.0 0.98 Aspen est.
(a) Recovered after separations.
(b) Phosphorus partitioning is not directly modeled in Aspen because of the small quantity,
most of which reports to the solid phase.
(c) Calculated using Boie’s equation (Boie 1953).
(d) The experimental system includes a continuous flow stirred-tank reactor (CSTR) followed
by a plug-flow reactor (PFR). The CSTR helps prevent overheating of the feed.
2.3.2 HTL Capital Costs
The HTL section capital costs are primarily based on estimates developed by Knorr et al. (2013) for a
wood HTL system under the National Advanced Biofuels Consortium (NABC), and scaled to the rate of
the sludge HTL plant. Major equipment costs adapted from the NABC work include the feed/product
exchangers, trim preheater, hot oil system, solids filter, and oil/water separator. Due to the corrosive
nature of the HTL environment, stainless steel 316-L is selected for process equipment The experimental
HTL system is constructed of stainless steel 316-L (with the exception of the CSTR, which is Inconel),
for which no measurable corrosion has been detected during testing of sludges. In addition, materials of
contruction testing, including systematic ex-situ coupon testing, in-situ coupon testing, and destructive
evaluation of system components indicated no measurable corrosion of stainless steel 316-L coupons
during 1000 hours exposure time. Feed/product heat exchange is modeled using two double-pipe type
exchangers in series. Attempts to obtain budgetary estimates for these exchangers were unsuccessful and
therefore capital costs are scaled on estimates provided by Knorr et al. (2013). The feed/product
14
exchanger areas are adjusted based on a more optimistic heat transfer rate that is thought to be attainable
for the goal case. The heat transfer coefficient is dependent on the physical properties of the sludge (e.g.,
viscosity, heat capacity, density) as well as flow properties (e.g., pipe diameter, line velocity). Testing of
the PNNL bench scale tube-in-tube heat exchanger has shown an overall heat transfer coefficient (U) of
50 Btu/hr/ft2/°F for heating a 15% wood slurry from 77°F to 549°F (25-287°C) with hot oil. The general
trend in viscosity among feedstock slurries at similar solids content has been observed in PNNL testing as
algae<sludge<wood and therefore a higher heat rate is expected for sludge compare to wood. In addition,
sludge viscosity should decrease with shear rate and consequently scaled up systems should provide for
higher heat rates. With higher line velocities and higher Reynolds Numbers, an overall U of 100
Btu/hr/ft2/°F for the combined feed/product exchangers is considered achievable in the goal case
timeframe. Further work is needed to characterize the rheology of the sludge at relevant line velocities
and temperature ranges, investigate configurations/strategies for improved sludge heating, evaluate the
effect of fouling on the heat transfer coefficient and ultimately demonstrate the achievable heat exchange
rate at larger than bench scale.
The HTL reactor is a jacketed serpentine pipe with heating oil in the annular space, similar to a
double-pipe heat exchanger. To minimize pressure drop, a configuration of two reactors in parallel is
assumed, similar to the design in Knorr et al. (2013). The developed parameters for each reactor are
listed in Table 6. ACCE was used to estimate the equipment cost of a fully jacketed pipe with elbow
fittings for the reactor. Based on the PNNL experimental runs to date, an oil jacket is recommended to
minimize heat losses over the length of the reactor. The reactor design in Knorr et al. (2013) also
assumed a serpentine pipe type reactor, but without an oil jacket. Therefore, the capital estimate
presented in this analysis is slightly more expensive than the Knorr design.
Table 6. HTL reactor specifications.
Spec Value Spec Value
Feed Rate, lb/hr (gal/hr) 18,300 (2,100) Overall Length, ft 530
LHSV, hr-1 6 Line velocity, ft/min 53
Reactor Volume, gal 350 Surface Area, ft2 550
Inside Pipe Diameter, inch 4 U, Btu/hr/ft2/°F 154 (Knorr et al. 2013)
Pipe Thickness, inch 0.67 Q, mmBtu/hr 2.4
Pipe Material SS316 Pipe section length, ft 25
Pipe Schedule XXH Number of pipe sections 21
Note: Two of the specified reactors are used in parallel for the process.
The capital costs for each of the major process components of the HTL section are given in Table 7.
Individual equipment costs as well as scaling assumptions and installation factors can be found in
Appendix B.
15
Table 7. Sludge HTL capital costs.
Item
Purchased,
million USD
(2014)
Installed,
million USD
(2014) Source
HTL Reactor System: Pumps, heat
integration, HTL reactor, knockout drums
$5.67 $11.34 Knorr et al. 2013; ACCE
(reactor)
Phase separation $0.47 $1.12 ACCE (product coolers);
Knorr et al. 2013
(separators)
Hot oil system for reactor and trim heater $0.48 $0.64 Knorr et al. 2013
Total $6.6 $13.1
2.4 HTL Aqueous Phase Treatment
2.4.1 Aqueous Treatment Design Basis
After separation from the biocrude product, the HTL aqueous phase is routed back to the WWTP’s
treatment train. Table 8 lists the measured HTL aqueous phase constituent concentrations from the
experimental run with Detroit sludge. While the organic and ammonia levels in the HTL water are high
relative to a typical WWTP influent, the overall volume is small. Table 9 provides a comparison of the
component loadings from the HTL water versus the average WWTP headworks (calculated for 100 MGD
at concentrations from Tchobanoglous et al. (2014). As shown in Table 9, the HTL water contributes 9%
and 18% of the BOD and ammonia load to the WWTP, respectively. The organic content, which is
essentially all dissolved, would likely be processed without issue by the activated sludge process (note
that the extra aerating power required is included in the analysis). The high ammonia, on the other hand,
could cause problems at some plants, such as toxicity impacts and/or pass through to the WWTP’s
discharge to the environment. Significantly high phosphate levels may also pose a problem. Some plants
may be able to recycle the HTL aqueous stream directly back to primary or secondary treatment without
pre-treatment. Others, particularly those that are already near their permitted effluent nitrogen and/or
phosphorus limit and/or do not already have an enhanced nutrient removal step in the main treatment
train, might have to treat this stream to reduce loads on the plant. In any case, it is anticipated that
effluent limits will become increasingly stringent and nutrient recovery will become more important for
WWTPs in the future. Therefore, as a conservative measure, treatment of the HTL water is included in
the design.
16
Table 8. HTL aqueous phase constituent analysis from experimental run
Constituent mg/L(a)
BOD(d) 22,900
COD 61,300
CODs Not measured
TOC (total organic carbon) 20,550
Total Carbon 24,700
TKN(b) (total Keldahl nitrogen) 7,670
TSS (total suspended solids) Not measured
Nitrogen (total)
Free ammonia-Nitrogen
7,700
4,900
Calcium <1
Phosphorus (total) 11
Potassium 280
Sodium 140
Sulfur 461
Silicon 294
Chloride 280
Nitrate 36
Phosphate <10
Sulfate 196
Alkalinity as CaCO3, total(b) 15,700
Dissolved Organics
Acetic Acid
Acetone
Ethanol
Methanol
P-Cresol
Propanoic acid
5,450
1,090
240
6,870
1,170
2,040
Phenolics, total recoverable(b) 69
pH 7.8
(a) Detroit sludge; WW-06 SS-2 (20% solids feed; LHSV=4 hr-1)
(b) Detroit sludge; WW-06 (composite of LHSV 2 hr-1 and 4 hr-1)
17
Table 9. Comparison of HTL effluent water and WWTP headworks component flows.
Untreated WWTP
Headworks
Wastewater
HTL Aqueous
Phase Effluent
Contribution of
HTL Effluent water
to Headworks
Loadings(%)
Flow Rate for 100 MGD WWTP, gal/min 69,444 67(a) 0.2%
Constituent lb/day(b) lb/day % Contribution
BOD(d) 166,907 16,512 9%
COD 423,945 44,200 9%
CODs 147,713 Not measured Not measured
TOC (total organic carbon) 136,864 14,817 10%
Total Carbon Not listed 17,810 Not listed
TKN(d) (total Keldahl nitrogen) 29,209 5,530 16%
TSS (total suspended solids) 162,735 Not measured Not measured
Nitrogen (total)
Free ammonia-Nitrogen
29,209
16,691
5,552
3,533
16%
17%
Calcium Not listed <0.7 Not listed
Phosphorus (total) 4,673 8 0.2%
Potassium 13,353 202 1.5%
Sodium Not listed 101 Not listed
Sulfur Not listed 332 Not listed
Silicon Not listed 212 Not listed
Chloride 49,238 202 0.4%
Nitrate 0 26 100%
Phosphate 653 <7 1%
Sulfate 30,043 141 0.5%
Alkalinity as CaCO3, total(d) Not listed 11,320 Not listed
(a) Using Table 8 data and assuming 100 dry ton/day sludge is generated
(b) Based on concentrations given in Tchobanoglous et al. 2014
Possible treatment options for ammonia removal from the HTL aqueous stream include physical
methods, such as ammonia stripping or membranes, as well as biological methods. An air stripping
system was selected for the design as this option is well established and reliable for lowering the
ammonia content of wastewater (EPA 2000). As shown in Figure 5, the HTL aqueous stream is initially
treated with lime (CaO) to raise the pH to 11 and shift the NH3/NH4+ equilibrium to the gas phase. Lime
treatment causes CaCO3 and CaPO4 to precipitate and is therefore effective for removing phosphate, if
present. After clarification of precipitated CaCO3 and CaPO4, the effluent water containing dissolved
ammonia gas is contacted with an air stream in a packed tower where the ammonia gas is stripped from
the wastewater. Preliminary modeling indicates that air stripping will also remove volatile oxygenates
(e.g., ethanol, acetone, acetic acid, butanol) from the wastewater at levels that are likely too high to
recover a marketable ammonia product. Therefore, the ammonia/air stream is treated with a thermal
oxidation unit (THROX) where ammonia and organics are catalytically combusted to N2, CO2 and H2O
and released to the atmosphere.
18
A membrane system and biological based methods were also examined as possible options for
ammonia removal from the HTL aqueous stream. It is possible a membrane system could be used to
recover saleable ammonia byproduct, but the efficacy of a membrane is highly uncertain as fouling from
precipitated salts (following pH adjustment) is a significant risk. Biological treatment systems for the
anaerobic oxidation of ammonium (anammox) in WWTP sidestreams have been installed at several
facilities in the U.S. and Europe WWTP (Pugh 2015). Preliminary discussions with vendors indicate that
a hybrid system consisting of an anaerobic blanket and anammox (Lu et al. 2016) is a plausible option.
However, this option does not allow for nutrient product recovery as ammonia is oxidized to nitrogen gas
in the biological process. Recovery of a nitrogen product pure enough to be saleable will likely require
removal of the organics prior to the ammonia recovery step. For example, catalytic hydrothermal
gasification (CHG), which can be used to convert organics in wastewater to methane (Jones et al. 2014),
could be coupled with air stripping or a membrane system to recover a pure ammonia byproduct.
AirStripper
Treated WaterTo Headworks
Lime softening &
clarifierHTL Aqueous
Quicklime (CaO)
CaPO4, CaCO3
Air
ThermalOxidizer
Natural gas
N2, H2O
Figure 5. Process flow for HTL aqueous phase treatment.
Table 10 summarizes some advantages and disadvantages of each treatment method considered for
this initial technology screening. Note that this list is not intended to be comprehensive, as there are
likely other possible treatment options that could not be included in this initial screening. The treatment
scheme using ammonia stripping is preliminary and as such, further experimental testing and analysis is
needed to test its feasibility and to determine the most appropriate and economical treatment option for
the HTL aqueous phase. Additional work is also needed to determine the overall impact of recycling the
HTL water on the WWTP. For example, there may be components other than ammonia that could build
up in the HTL water and at the least, require a purge.
19
Table 10. Possible HTL water ammonia removal/recovery options.
Option Pros Cons
No Treatment (direct recycle
to WWTP)
Little/no cost May cause WWTP effluent nitrogen level to
increase.
NH3 Stripping Conventional technology;
Potential recovery of NH3
byproduct
Organics in NH3 stream may be too high for saleable
product; destruction of NH3 is expensive and GHG-
intensive; Does not remove other nitrogen.
Membrane Off-the-shelf; Potential
recovery of NH3 byproduct;
Simple system; small
footprint
Potential membrane fouling with organics and/or
suspended solids; Organics in NH3 stream may be
too high for saleable product; Does not remove other
nitrogen.
Biological (Anammox or
Nitrification-Denitrification)
Fairly well-established for
WWTP AD centrate
treatment
Additional biological operations and associated
potential for upsets; May require methanol (for
nitrification-denitrification; No option for product
recovery (produces N2 gas).
2.4.2 Aqueous Treatment Capital Costs
A separate, preliminary model for the aqueous phase treatment section was developed with Aspen’s
electrolyte package to enable more accurate prediction of lime requirements for pH adjustment and get
initial estimates for capital and operating costs. Capital costs for the lime softening, air stripping, and
thermal oxidation equipment were estimated with ACCE. Table 11 shows the purchased and installed
cost for the complete treatment process.
Table 11. HTL aqueous phase ammonia removal system capital costs
Equipment
Purchased, million
USD (2014)
Installed, million
USD (2014) Source
Lime softening and clarification $0.43 $1.04 ACCE, v. 8.8
Air stripping $0.63 $1.51 ACCE v. 8.8
Thermal oxidation $0.22 $0.53 ACCE, v. 8.8
Total $1.3 $3.1
2.5 Centralized Biocrude Upgrading
The HTL biocrude requires catalytic processing to remove oxygen, nitrogen, and sulfur before it can
be used as fuel blendstock. Biocrude produced from multiple plants within a geographical area is
collected at a central facility where it is hydroprocessed and finished into fuel products. The upgrading
plant is assumed to process 2,700 BPSD of biocrude, which is equivalent to the output of ten HTL
facilities, each processing 110 dry ton/day sludge. The upgrading plant flow diagram is shown in Figure
6 and based on that modeled in Jones et al. (2014). Prior to upgrading, a desalting process similar to that
used in a petroleum refinery is used to remove inorganic components from the biocrude to prevent
deactivation of the hydrotreating catalyst. The biocrude is then pumped to ~1500 psia, mixed with
compressed hydrogen, and preheated to the hydrotreater reactor temperature of ~750°F (400°C).
Hydrogen is produced onsite via steam reforming of the process off-gas and additional purchased natural
20
gas (see Section 2.5.3). A guard bed directly upstream of the hydrotreater is used to remove soluble iron
that may be present in the biocrude. In contrast to algae HTL biocrude (Jarvis et al. 2016), no solubilized
iron complexes were observed in the sludge HTL biocrude and therefore a filter may be sufficient for
mineral removal in lieu of a guard bed in the future. However, there may be other mineral compounds
present that are unlikely to be captured by a filter. Further investigation is needed in this area to verify
effective biocrude cleanup strategies. During the hydrotreating process, biocrude oxygen is converted to
CO2 and water, nitrogen is converted to ammonia, and sulfur is converted to hydrogen sulfide. The
resulting hydrocarbon product consists of a mixture of paraffins, olefins, naphthenes and aromatics that
lie within the gasoline, jet and diesel boiling ranges. The hydrotreater reactor effluent is cooled to
condense the produced water and hydrocarbons and the organic phase is fractionated into four boiling
point cuts: C4 minus, naphtha range, diesel range and heavy oil range material. The heavy cut is cracked
in a conventional hydrocracker to produce additional naphtha and diesel range products. As product cut
analysis is not yet available, simulated distillation data are used to estimate the volumes of naphtha, diesel
and heavy oil boiling range material. The wastewater from the process is high in ammonia and therefore
will likely require treatment prior to sending to a WWTP. A stripping system similar to that described for
HTL water treatment is assumed for the upgrading plant.
Biocrude Feed
HydrotreaterR-310
P-310
H-311
Makeup CompressorC-311
Feed/Product HX-318A
Air Fin CondenserH-316
HP FlashV-311
LP FlashV-315
PSAV-300
Offgas to A350
To Water Treatment
Lights ColumnT -310
Diesel ColumnT-330
To Gasoline Pool
To Diesel Pool
To A350 Hydrocracker
E-310
Naphtha ColumnT -320
336
Hydrogen From A400
Recycle CompressorC-310
HX-320
Figure 6. Process flow for HTL biocrude hydrotreating.
2.5.1 Biocrude Hydroprocessing Design Basis
The design basis for hydrotreating is presented in Table 12. The experimental data was collected
from bench scale flow reactors at PNNL. The modeled goal case assumes improved hydrotreated product
yield and reactor space velocity relative to that demonstrated in the testing. Based on the experience of
21
the experimental team, these values are expected to be attainable within the time frame of the goal case
with further research and development.
Table 12. Sludge biocrude hydrotreating experimental results and model assumptions.
Component
Experimental
(HT-62005-60)
Model
(goal case)
Temperature, °F (°C) 752 (400) 752 (400)
Pressure, psia 1540 1515
Catalyst
Sulfided?
CoMo/alumina-F
Yes
CoMo/alumina
Purchased presulfided
LHSV, vol./hour per vol. catalyst 0.18 0.50
WHSV, wt./hr per wt. catalyst 0.29 0.81
HTL biocrude feed rate, ml/h 5.6 Commercial scale
Total continuous run time, hours 302 (total run) Not applicable
Chemical H2 consumption, wt/wt HTL biocrude (wet) 0.046 0.044
Product yields, lb/lb dry biocrude (vol/vol wet biocrude)
Hydrotreated oila
Aqueous phase
Gas
0.82 (0.99)
0.14 (0.13)
0.08
0.84 (0.97)
0.13 (0.19)
0.07
Product oil, wt%
C
H
O
N
S
85.6%
14.6%
1.0%
<0.05%
7-10 ppm
85.3%
14.1%
0.6%
0.04%
0.0%
Aqueous carbon, wt% 0.10% 0.17%
Gas analysis, volume%
CO2, CO
CH4
C2+
NH3
NH4HS
0%
51%
49%
Not measured
Not measured
0%
33%
38%
26%
3%
TAN, feed (product) 59 (<0.01) Not calculated
Viscosity@40 °C, cSt,
feed (product)
400 (2.7)
Not calculated
Density@40 °C, g/ml,
feed (product)
0.98 (0.79)
0.98 (0.79)
(a) Yield after phase separation
The hydrocarbon compounds used to model the hydrotreated product are similar to those used for
previous algae HTL modeling (Jones et al. 2014) and are listed in Appendix C. Figure 7 shows the
boiling point curves from simulated distillation (ASTM Method D2887) for the hydrotreated product
from experimental testing and the modeled hydrotreated product. Actual distillation was also performed
on the hydrotreated product and matched well with the D2887 results.
22
Figure 7. Boiling point distribution (ASTM D2887) for hydrotreated product from sludge biocrude.
The heavy oil fraction from hydrotreating is assumed to be hydrocracked into additional gasoline and
diesel range fuel. No experimental data are yet available for hydrocracking of hydrotreated HTL
biocrude, so it is assumed that the heavy cut is processed similar to petroleum operations. The
hydrocracking assumptions for the model are given in Table 13.
Table 13. Hydrocracking model assumptions.
Process Basis Assumptions
Hydrocracking heavier than
diesel portion of hydrotreated
HTL biocrude
No experimental data, assumed to be
similar to conventional hydrocrackers
LHSV=1
Temperature: 734 °F (390 °C)
Pressure: 1035 psia
H2 chemical consumption:
0.004 wt/wt heavy oil
Product breakdown:
Gas (excluding excess H2); 3 wt%
Liquid fuels: 96 wt%
Aqueous: 1 wt%
2.5.2 Biocrude Hydroprocessing Capital Costs
Table 14 shows the capital costs for the hydrotreater and hydrocracker systems, which are based on a
review of conventional naphtha, diesel and kerosene hydrotreaters as published in the IHS 2014 Yearbook
and from Petroleum Refining Technology and Economics (Gary 2007). Additional cost details can be
found in Appendix B.
23
Table 14. Hydrotreater and hydrocracker capital costs.
Item or Area
Purchased, million
USD (2014)
Installed, million USD
(2014) Source
Hydrotreater system (2732 BPSD feed) 21.1 33.6 IHS 2014a
Hydrocracker system (1020 BPSD feed) 4.4 6.6 IHS 2014a
The costs spanned a broad range from simple naphtha hydrodesulfurization units, to multi-stage
hydrocrackers. In addition to the reactor(s), each system at least includes recycle compressors, multi-
stage flash systems and distillation. The hydrotreater system also includes the cost of a desalter and guard
bed. Costs for a generic hydrocracking system (2000 psia) were chosen as the basis for hydrotreating and
hydrocracking. While these costs are generally applicable to HTL oil hydrotreating as they employ
similar temperatures and pressures, conventional refining space velocities tend to be higher. Reactor cost
sensitivity is considered in Section 4.0.
2.5.3 Hydrogen Generation Design and Cost Basis
The hydrogen plant is a conventional natural gas based steam reformer. Most of the off-gas is used to
fire the reformer. However, a portion of the off-gas is compressed and mixed with makeup natural gas
which is then sent to a hydrodesulfurizer (HDS) unit. Figure 8 (Jones et al. 2014) shows the simplified
flow scheme for hydrogen generation by steam reforming of natural gas (IHS 2014b; Meyers 2004; H2A
2013) combined with the off-gas streams from hydrotreating and hydrocracking. Hydrogen for the HDS
unit is supplied by the off-gas stream. The gas exiting the HDS unit is then mixed with superheated
steam and sent through an adiabatic pre-reformer to convert C2+ compounds to methane prior to entering
the main steam reformer to produce syngas. This reduces the rate of coking in the main reformer. The
syngas hydrogen content is increased by high temperature shift. After condensing out the water, the
hydrogen is purified by pressure swing adsorption (PSA). Off-gas from the PSA is recycled to the
reformer burners.
24
Figure 8. Process flow for hydrogen production (from Jones et al. 2014).
Saturated and superheated steams are generated by recuperating heat from the reformer exhaust and
cooling the product from the water gas shift reactor. The generated steam is used in the reformer and also
to provide process heat, including the distillation column reboilers.
The design assumptions are shown in Table 15. Gibbs minimization reactors are used to model the pre-
reforming, methane reforming and burner reactions. The reactor methane conversion of 80 mole %
matches that reported by IHS (2014b).
Table 15. Hydrogen plant model assumptions.
Equipment Assumptions
Pre-reformer
Outlet temperature 925 F (496 C)
Outlet pressure 429 psia
Steam/carbon ratio 3.5
Methane Reformer
Steam pressure 670 psia
Outlet temperature 1562 F (850 C)
Outlet pressure 399 psia
Burners Bridge wall temperature 1800 F (982 C)
Pressure Slightly positive
Shift Reactor
Outlet temperature 568 F (300 C)
Outlet pressure 388 psia
Approach to equilibrium 98%
PSA Hydrogen delivery pressure 376 psia
Hydrogen recovery 90%
Capital costs for hydrogen generation are taken from the IHS PEP 2014 Yearbook and scaled to the
necessary hydrogen production rate using the IHS scale factor. The equipment includes a sulfur guard
PSA
Shift
Steam Drum
Methane Reformer
Hydrogen
Offgas
Air
Offgas
Natural Gas
Flue Gas
Condensed Water
HDS
Pre-Reformer
SuperheatedSteam
25
bed, pre-reformer, primary reformer with nickel catalyst, high temperature shift reactor, pressure swing
adsorption unit, waste heat recovery producing high pressure steam and all associated outside battery
limit equipment. Conventional steam reformer hydrogen plants range in scale from 1 to 100 million scf
of hydrogen per day. The hydrogen plant scale needed for the centralized biocrude upgrading plant is at
the low end: 9 million scf per day. The purchased capital and installed capital cost for this plant is $14.2
million and $27.2 million (in 2014 dollars), respectively. Additional cost details can be found in
Appendix B.
3.0 Process Economics
The process economic analysis involves first determining the total capital investment (TCI), the
variable operating costs, and the fixed operating costs for each of the HTL and biocrude upgrading plants.
Discounted cash flow rate of return analysis is then used to determine the fuel production cost using
standard methodology used for all BETO design cases. The summary economics and performance for the
HTL plant and the upgrading plant are presented in the following sections. Sensitivity analysis around
key technical and economic assumptions is presented in Section 4.0.
3.1 Sludge HTL Plant
Table 16 lists the feed and biocrude product flow rates for the sludge HTL plant goal case.
Table 16. Annual feed and biocrude product rates for the sludge HTL plant.
Stream Million gallons/year Million lb/year
Dry sludge (15% ash) feed N/A 73
Dry, ash-free sludge feed N/A 62
Total slurry feed (25% solids) 33 290
HTL biocrude product 3.8 30
HTL aqueous phase 25 9.4
HTL solids (dry) N/A 2000
3.1.1 Total Capital Investment
Table 17 summarizes the costs presented in Section 2.0 for the sludge HTL plant goal case including
the balance of plant items such as the tank farm, flare and cooling water system. The HTL section of the
plant contributes 72% the total installed capital.
26
Table 17. Total capital investment for sludge HTL plant.
Cost Item Million US Dollars (2014$)
Sludge dewatering 1.43
HTL biocrude production 13.1
HTL aqueous phase treatment 3.09
Balance of plant 0.62
Total Installed Cost (TIC) 18.2
Buildings (4% of TIC) 0.73
Site development (10% of TIC) 1.82
Additional piping (4.5% of TIC) 0.73
Total Direct Costs (TDC) 21.5
Indirect Costs
Prorated expenses (10% TDC) 2.15
Home office & construction fees (20% TDC) 4.30
Field expenses (10% TDC) 2.15
Project contingency (10% TDC) 2.15
Startup and permits (10% TDC) 2.15
Total Indirect 12.9
Fixed Capital Investment (FCI) 34.4
Working Capital (5% of FCI) 1.72
Land(a) 0.009
Total Capital Investment (TCI) 36.2
(a) Scaled on Dutta et al. 2011
3.1.2 Operating Costs
Variable operating costs and supporting assumptions for the sludge HTL plant are given in Table 18.
The largest contributors to annual operating cost are quicklime for HTL aqueous phase ammonia
stripping, natural gas for the HTL process, extra aeration power needed for processing the recycled HTL
water in the WWTP’s biological treatment step, and polymer consumed for sludge dewatering.
27
Table 18. HTL plant variable operating costs.
Variable Value Source
Total Cost (2014),
million USD/year
Sludge Dewatering
Polymer, $/lb (2013$) 1.73 City of Detroit 2014 0.33(a)
HTL Processing
Natural Gas, $/1000 scf (2014$) 5.62 EIA 2014a 0.39
Electricity, ¢/kWh (2014$) 7.09 EIA 2014b 0.10
HTL Aqueous Phase Treatment
Quicklime, $/ton (2014$) 107 USGS 2016 0.42
Natural gas (for THROX unit), $/1000 scf
(2014$)
5.62 EIA 2014a 0.18
Extra Aeration at WWTP
Electricity, kWh/lb COD removed
(assuming100% COD removal)
0.40 Wan et al. 2016 0.36
Total 1.78
(a) For a polymer dose of 10.5 lb/dry ton secondary sludge.
Fixed costs for the HTL plant are shown in Table 19. Salaries are taken from Dutta et al. (2011) and
converted to a 2014 dollar basis using US Bureau of Labor Statistics labor cost indices. The factors for
benefits and maintenance, insurance and taxes are the standard assumptions used for BETO design cases.
Table 19. HTL plant fixed operating costs.
Position Title Number
Total Cost (2014),
million USD/year
Conversion Plant (unburdened)
Plant Manager 1 0.15
Plant Engineer 1 0.07
Maintenance Super 1 0.06
Lab Manager 1 0.06
Shift Supervisor 3 0.14
Lab Technician 1 0.04
Maintenance Tech 1 0.04
Shift Operators 4 0.19
Yard Employees 1 0.03
Clerks & Secretaries 1 0.04
Subtotal 0.81
Overhead & maintenance 90% of labor & supervision 0.73
Maintenance capital 3% of TIC 0.98
Insurance and taxes 0.7% of FCI 0.24
Total Other Fixed Costs 2.76
28
3.1.3 Minimum Biocrude Selling Price
The minimum fuel selling price (MFSP) of the biocrude is determined using a discounted cash flow
rate of return analysis. The MFSP is the plant gate selling price of the fuel product that makes the net
present value of the project equal to zero with a 10% discounted cash flow rate of return over a 30 year
plant life and 40% equity with the remainder debt financed at 8% interest for a 10 year term (see Table 1,
Section 1.3). The resulting MFSP for the 100 dry ton/day sludge HTL plant is $2.35/gge biocrude, or
$2.53/gallon. The modeled biocrude has a lower heating value (LHV) of 124,990 Btu/gal. A LHV of
116,090 Btu/gal (ANL 2016) is used to convert the heat value of the fuel products to a gasoline gallon
equivalent basis. Table 20 shows the breakdown of costs contributing to the MFSP of the biocrude. This
MFSP corresponds to a revenue from biocrude sales of $9.6 million/year. The payback period of the
project, calculated as the fixed capital investment divided by the average annual cash flow (Peters et al.
2004), is 13 years at the assumed 10% internal rate of return (IRR). This payback period does not
consider any avoided sludge disposal costs that would be incurred due to reduction of solids from HTL.
Figure 9 shows the contribution of each section of the HTL plant to the overall MFSP. The HTL section
of the plant constitutes 72% of the production cost. Sensitivity analysis around key financial and
technical assumptions and their impact on the MFSP is presented in Section 4.0.
Table 20. Biocrude MFSP cost breakdown.
$/gge
biocrude
$/year
(2014 USD)
Natural Gas 0.14 $600,000
Quicklime, Polymer 0.18 $800,000
Electricity 0.11 $500,000
Fixed Costs 0.68 $2,800,000
Capital Depreciation 0.27 $1,100,000
Average Income Tax 0.16 $700,000
Average Return on Investment 0.81 $3,300,000
Total 2.35
29
Figure 9. Sludge HTL biocrude production cost for goal case.
3.2 Biocrude Upgrading Plant
Table 21 lists the feed and product flow rates for the centralized sludge biocrude upgrading facility.
The plant processes 2,730 BPSD of biocrude feed (from 10 HTL plants) and produces 2,660 BPSD of
diesel and gasoline blendstocks.
Table 21. Annual feed and product flows for centralized biocrude upgrading plant (processing biocrude
from 10 HTL plants).
Stream Million gal/year Million lbs/year
Biocrude feed 37.9 296.2
Hydrotreated oil 36.6 242.6
Diesel blendstock 27.6 (29.6 million gge/year)(a) 183.8
Naphtha (gasoline blendstock) 9.2 (9.1 million gge/year) (a) 56.5
(a) Based on a gasoline LHV of 116,090 Btu/gal (ANL 2016).
3.2.1 Total Capital Investment
Table 22 summarizes the costs presented in Section 2.0 for the sludge HTL biocrude upgrading plant
including the steam cycle and balance of plant items such as the tank farm, flare and cooling water
system. The hydrotreating section of the plant has the single highest installed capital cost.
30
Table 22. Total capital investment for biocrude upgrading plant.
Cost Item Million US Dollars (2014$)
Hydrotreating 33.6
Hydrocracking 6.6
Hydrogen plant 27.2
Steam cycle 1.6
Balance of plant 6.5
Total Installed Cost (TIC) 75.6
Buildings (4% of TIC) 3.0
Site development (10% of TIC) 7.6
Additional piping (4.5% of TIC) 1.8
Total Direct Costs (TDC) 87.9
Indirect Costs
Prorated expenses (10% TDC) 8.8
Home office & construction fees (20% TDC) 17.6
Field expenses (10% TDC) 8.8
Project contingency (10% TDC) 8.8
Startup and permits (10% TDC) 8.8
Total Indirect 52.8
Fixed Capital Investment (FCI) 140.7
Working Capital (5% of FCI) 7.0
Land (6% of TPEC)(a) 2.7
Total Capital Investment (TCI) 150.4
(a) Total purchased equipment cost
31
3.2.2 Operating Costs
Table 23 and Table 24 list the plant variable and fixed operating costs and supporting assumptions for
the biocrude upgrading facility. Biocrude cost is the large majority (95%) of the operating cost and
therefore efforts aimed at reducing its production cost are critical to reducing the final fuel MFSP.
Table 23. Biocrude upgrading plant variable operating costs.
Variable Value Source
Total Cost
(2014), million
USD/year
Biocrude cost, $/gge (2014$) 2.42(a) This analysis 99.9
Hydrotreating catalyst, $/lb (2014$)
(2 year life)
16.6 IHS 2014c 0.40
Hydrocracking catalyst, $/lb (2014$)
(5 year life)
16.6 IHS 2014c 0.03
Hydrogen plant catalyst, $/1000scf H2
(2014$) (5 year life)
0.0205 IHS 2014b 0.06
Natural gas, $/1000scf (2014$) 5.62 EIA 2014a 2.79
Cooling tower chemical, $/lb (2007$) 1.36 Humbird et al. 2011 0.01
Boiler chemical, $/lb (2007$) 2.27 Humbird et al. 2011 0.01
Electricity, ¢/kWh (2014$) 7.09 EIA 2014b 0.92
Water makeup, $/ton (2001$) 0.20 Humbird et al. 2011 0.05
Wastewater fee, $/ton (2001$) 0.48 Peters et al. 2004 0.07
Total 104.2
(a) Includes 10 cents/gge cost for biocrude transportation to the refinery (Sheppard 2011).
Table 24. Biocrude upgrading plant fixed operating costs.
Position Title Number
Total Cost (2014),
million USD/year
Conversion Plant (unburdened)
Plant Manager 1 0.15
Plant Engineer 1 0.07
Maintenance Super 1 0.06
Lab Manager 1 0.06
Shift Supervisor 3 0.14
Lab Technician 3 0.12
Maintenance Tech 6 0.24
Shift Operators 25 1.20
Yard Employees 4 0.11
Clerks & Secretaries 1 0.04
Subtotal 2.19
Overhead & maintenance 90% of labor & supervision 1.97
Maintenance capital 3% of TIC 4.05
Insurance and taxes 0.7% of FCI 0.99
Total Other Fixed Costs 9.18
32
3.2.3 Minimum Fuel Selling Price
The MFSP for the goal case centralized upgrading plant is $3.46/gge fuel blendstock. Table 25
shows the breakdown of costs contributing to the fuel blendstock MFSP. Note that the cost of the
biocrude feed is given per gge of upgraded fuel, considering that 1.05 gge of biocrude is needed to make
1 gge of upgraded fuel. Figure 10 shows the contribution of each section of the upgrading plant to the
overall MFSP. Again, the overriding impact of biocrude price on the fuel production price is evident.
Sensitivity analysis around key financial and technical assumptions and their impact on the MFSP is
presented in Section 4.0.
Table 25. Upgraded fuel blendstock MFSP cost breakdown.
$/gge
final
fuel $/year
Biocrude Feed 2.58 $99,900,000
Natural Gas 0.07 $2,800,000
Catalysts & Chemicals 0.01 $500,000
Waste Disposal 0.002 $100,000
Electricity and other utilities 0.02 $1,000,000
Fixed Costs 0.24 $9,200,000
Capital Depreciation 0.002 $4,700,000
Average Income Tax 0.07 $2,800,000
Average Return on Investment 0.46 $17,700,000
Total 3.46
Figure 10. Upgraded fuel production cost for the goal case.
33
4.0 Economic and Technical Sensitivities
4.1 Sensitivity Analysis
Figure 11 shows the results of sensitivity analysis around the key technical and economic modeling
assumptions. A wide range of plant scale of 50 to 950 dry ton/day was selected to include the largest
plants in the country. Collectively this size range covers 52% of the total WWTP capacity represented in
the 2012 CWNS data (EPA 2016). Smaller sized plants are not economically feasible under conventional
equipment scaling assumptions for stick-built plants (i.e., the sixth-tenths rule), however biocrude
collection scenarios centered around large WWTPs in densely populated regions could include smaller
plants while still meeting the goal fuel MFSP, as explored in Section 4.2. As sludge management and
disposal currently constitutes a major expense to the wastewater industry, the impact of avoided sludge
disposal cost (feedstock cost) is an important aspect of reflecting the current reality for WWTPs. In some
cases, the HTL plant would be owned by the WWTP, in others by a private company. In the latter case,
the avoided sludge disposal cost may be in the form of a fee that WWTPs would pay to the HTL
owner/operator that would allow sharing of the cost savings between the two parties. A range of $200/dry
ton credit (avoided cost) to a $25/dry ton sludge cost was assessed for the sensitivity analysis. Biocrude
yield and feed solids content were varied based on minimum values observed in experimental testing and
maximum values thought to be possible with extensive research and development. The biocrude yield
will vary depending on the composition of the sludge, generally following the trend of
lipids>proteins>carbohydrates (Biller and Ross 2011). Hence, blending of sludge with high-lipid waste
feedstocks such as FOG from WWTP skimming operations and animal rendering wastes could result in
significantly higher yields. Note that there may be competition for these feedstocks in certain areas and
this aspect needs to be considered on an individual plant basis when determining project feasibility.
Sludge ash content will also vary considerably from plant to plant. As a comparative example, ash
content of sludges provided by MetroVancouver in the LIFT study ranged from 7 to 12 wt% (Marrone
2016), while the mixed sludge provided by the City of Detroit was 26 wt%. We believe the Detroit
sludge to be on the upper end and MetroVancouver to be on the lower end of the spectrum and
accordingly, a range of 5-30% ash content was investigated. Note that the ash sensitivity only considers
changes in yield that are proportional to organic content and does not include possible influences of ash
content on HTL chemistry (e.g., catalytic) or oil separations (e.g., biocrude adhering to solids). There is
uncertainty regarding the amount of extra heat exchange capacity that will be needed due to fouling of the
sludge preheaters. The overall heat exchanger coefficient (U) was varied by -50% to + 25% of the base
case value to illustrate the potential impact of this factor on cost.
As shown in Figure 12, variability in the plant scale has the greatest impact on MFSP. Biocrude
yield, sludge ash content and the overall heat transfer coefficient for the feed/product heat exchangers are
key technical parameters that significantly affect production cost as well. Consideration of avoided
sludge disposal cost, uncertainty in HTL capital cost, and the assumed IRR used in the discounted cash
flow calculations are key economic assumptions impacting the MFSP. A 10% IRR is assumed for all
conversion pathway TEAs in BETO’s portfolio, however WWTP/HTL plant owners may want a higher
project IRR.
34
Figure 11. Sensitivity analysis for sludge HTL plant.
Figure 11 shows the sensitivity of final upgraded fuel MFSP to several economic and technical
parameters for the biocrude upgrading plant model. Biocrude price was varied widely according to the
approximate range seen from the variable HTL plant scale in Figure 11. Upgrader plant scale was varied
from 1,000 to 5,000 BPSD final fuel. For plant scales above the base case, it is likely that a biocrude
supply draw radius larger than the base case of 100 miles would be necessary. A sensitivity case where
the biocrude is upgraded at an existing petroleum refinery (TCI=0) is also considered. As shown, this
option could potentially reduce the MFSP by $0.50/gge. However, more work is needed to characterize
the processing needs of biocrude and the oxygen, nitrogen and sulfur limits that could be tolerated in a
petroleum refinery. The hydrotreated oil yield range shown represents the minimum observed in the
laboratory and the maximum thought to be achievable with the highest quality biocrude. It may be
possible to eliminate the hydrocracker through the use of a specially designed hydrotreating catalyst that
also has cracking capabilities. In this case, the heavy residual would be recycled to the reactor with a
small purge. This case is also included as a sensitivity. The sensitivity analysis shows that biocrude price
has the greatest impact on upgrading production cost. Plant scale, upgrading at an existing refinery, and
hydrotreated oil yield significantly affect cost as well. It is important to note that inherent to the success
of a centralized biocrude upgrader scenario are the contractual arrangements and logistical coordination
that would be required between owners/operators of the HTL/WWTP plants and the regional upgrading
plant.
35
Figure 12. Sensitivity analysis for biocrude upgrading plant.
4.2 Regional Resource and Fuel Potential Analysis
As shown in the sensitivity analysis, plant scale is a key cost driver for the sludge biocrude
production cost. The base case of 110 dry ton/day plant is approximately the minimum HTL plant scale
that is feasible within BETO’s $3/gge target when only considering a single WWTP size feeding an
upgrader of 2,700 BPSD fuel production capacity. However, WWTPs larger than 110 dry ton/day can
produce biocrude cheaper than the base case price of $2.35/gge. In this way, blending biocrude from
smaller WWTPs with biocrude from the largest WWTPs could maximize fuel production while still
meeting a feasible average biocrude price for the upgrader.
Using the sludge production dataset published by Seiple et al. (2017) and the HTL and biocrude
upgrading plant cost production models, a siting analysis was performed to identify specific regions of the
country that could support upgraded fuel production at or below $3.50/gge. This analysis is intended to
provide an initial estimate of the general regions of interest rather than identify specific siting locations.
Siting was constrained to the 15,014 existing WWTPs ranging from 0.001 to 812 MGD influent flow.
This was done in part to simplify the siting model, but also because WWTPs are highly spatially
distributed in the U.S. and large scale upgrader facilities are likely to be proximal to biocrude sources due
to transport costs and logistics. It is assumed that each WWTP has its own HTL plant and could send its
biocrude to a central upgrader site. An initial pool of 2,621 siting candidates was developed by
identifying all WWTPs capable of producing $3.50/gge fuel when aggregating biocrude within a travel
radius of 100 miles, without considering resource competition from neighboring facilities. Although only
candidate WWTPs could represent possible upgrader locations, resources from all WWTPs were
considered during siting. Siting begins at the candidate upgrading plant location with the highest total
36
fuel blendstock production potential. Resource competition is simulated by ensuring that each WWTP
can only participate in a single service territory. Within a given service territory, biocrude is
incrementally aggregated by prioritizing neighbors by total biocrude rather than transport distance.
Aggregation continues until the fuel price limit can no longer be achieved. Unused neighbors remain
available for subsequent siting iterations, allowing service territories to overlap but ensuring resource
pools are mutually exclusive. After each siting, the total fuel production potential is re-calculated for
remaining candidates using any remaining neighbors. As siting progresses, some viable candidates
become non-viable as their neighbors are scavenged by higher priority candidates. Siting continues until
all candidates are either sited or no longer able to achieve the fuel price limit.
The results of the upgrading plant siting analysis are presented in Figure 13. A total of seven
centralized upgrading sites are feasible in the U.S. given a fuel price limit of $3.50/gge. In total, these
sites utilize 34% (13,000 dry ton/day) of the total daily sludge feedstock in the U.S. to generate
approximately 1.32 million gallons/day of fuel blendstock. A total of 295 WWTPs (2%) were utilized
during siting. They ranged in size from 0.05 to 812 MGD wastewater influent flow, with an average of
40 MGD. Only two facilities had a flow rate less than 1 MGD. On average, candidate sites utilized 13%
(range of 3-31%) of their available neighbors, but consumed 72% (range of 29-94%) of available
biocrude produced within the service area. Regional biocrude aggregation increased fuel blendstock
production output by 225%, over the 0.41 million gallons/day of fuel that could be produced at 9 WWTPs
capable of independently producing fuel at or below $3.50/gge.
Figure 13. Regional fuel production at $3.50/gge from sludge HTL and biocrude upgrading.
37
5.0 Conclusions and Recommendations
This analysis shows that fuel blendstocks generated from HTL of sludge and centralized biocrude
upgrading have the potential to be competitive with fossil fuels. The estimated plant gate MFSP of
$3.46/gge final fuel blendstock for the goal case conceptual design for 2022 is within the tolerance of
BETO’s $3/gge programmatic goal for biofuel production cost. This analysis illustrates the feasibility of
HTL for point-of-generation conversion of waste feedstock at a scale 1/20th that of the standard
lignocellulosic biorefinery scale typically used in BETO design cases. The relevance of this work reaches
beyond wastewater treatment sludge to lay the groundwork for application to other distributed wet wastes
and blends that together represent a significant resource of underutilized biomass.
Future research in several key areas is needed to support the goal case design, including:
Increased biocrude yields through
– Improved continuous solid/liquid separations and liquid/liquid separations. Research is needed to
investigate and test improved methods for enhanced phase separation such as coalescers,
continuous solids separation, solids washing, feed de-ashing, and other approaches.
– Incorporation of FOG and other fat-rich waste resources, such as skimmed FOG from the
WWTP, collected restaurant grease trap waste, and regional rendering. Further work is necessary
to determine and test appropriate sludge/FOG blends.
– Increased feed solids concentration. Concentrating the solids increases the amount of carbon
reporting to the biocrude phase during HTL and improves economics through decreased capital
costs and increased energy efficiency. Further work is needed to assess high pressure pumping
and HTL of sludge feed at 25% solids at various ash contents.
– Decreased ash/grit content in the sludge solids through WWTP operations such as improved
degritting and/or deashing technology, and reduced inorganic flocculating agent usage.
Strategies for more effective and scalable heat transfer at HTL operating pressure:
– Improved understanding of fundamental heat and mass transfer limitations for sludge heating and
associated improvements to the design for application at scale. This includes heat transfer studies
to better characterize the preheater equipment needs and rheology studies to determine viscosity
profile as a function of temperature and flow rate.
– Investigation of alternative heat exchanger designs for reduced material costs such as static mixer
heat exchangers, dual-shell configurations (Fassbender 1992), and staged heating with shell/tube
units for the latter portion of preheating.
Cost-effective treatment technologies for the HTL aqueous phase. Baseline testing is needed to
determine the effectiveness of various options (e.g., ammonia stripping, biological methods and
membranes) for nitrogen removal. Further investigation of conversion (e.g., catalytic gasification)
and treatment technologies through collaborative experimental and modeling work with universities
and industry is needed to understand the impacts of full integration with the WWTP.
Address scale challenges for distributed waste feedstocks through:
– Integration of additional waste feedstocks on a regional basis. Resource analysis will help to
identify available waste feedstock blends including manure, FOG, pulp sludge, food waste, etc…
38
Testing of different blends to establish overall yields and impact on quality is needed. Dry waste
may also be beneficial as it is cheap to transport relative to wet materials and could reduce
dewatering energy.
– Investigation of the feasibility of transporting sludge to enable larger scale HTL plants.
– Investigation of modular HTL systems for smaller scale applications.
Improved performance of biocrude upgrading through (Jones et al. 2014):
– Improved catalyst performance. Hydrotreating catalyst maintenance and stability are unknown,
as are regeneration protocols and lifetimes. Longer-term testing with biocrude and detailed
characterization of catalyst performance and deactivation modes are needed. Pretreatment steps,
such as desalting, need to be demonstrated.
– Development of hydrodeoxygenation and hydrodenitrogenation reaction kinetics to assist reactor
designs and better inform the choice of co-processing in a petroleum refinery. A possible method
to be more generally applicable might be to establish rate laws for sludge components such as,
lipids, carbohydrates and proteins.
– Both standard and comprehensive characterization of the major distillation fractions, gasoline
range, diesel range, and gas oil range for the HTL oil and the hydrotreated oil. The jet fuel range
should also be characterized, and an understanding of how to produce a jet cut without degrading
naphtha and diesel properties would also be useful. Testing for key final fuel qualities, such as
flash, octane, cetane, and cold flow properties is desirable.
– Demonstration of hydrocracking yields of the gas oil fraction. Combined hydrotreating-
hydrocracking is a possible way to reduce cost and could be achieved through catalyst research
and development.
– Investigation of compatibility of biocrude with petroleum refinery insertion points and other
coprocessing strategies.
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A.1
- Heat and Material Balances
A.2
109
128B
111
128
QH100
115
131
130
132
Feeders/PumpsP-101b
126
F-105Filter
AqueousTo A200
HTL Oil to Upgrading Plant
HTL Solids
105
121
135
R-101 HTL Plug Flow Reactors with Hot
Oil Jacket
A100 HYDROTHERMAL LIQUEFACTION
HTL Plant
Heat Exchangers
HX-100
Air Fin CondenserHX-102
129
Fuel Gas to Combustor
Three-phaseseparator
V-110
HeaterH-105
Heat from Hot OilA110
Heat from Hot OilA110
Feeders/PumpsP-101a
Dewatered Sludge(25% solids)
101
HeaterH-100
Heat from Hot OilA110
109
107
128
111 115
140
132
131
130
136
H-110
A.3
101 105 107 109 111 115 121 126 128 129 130 131 132 135 136 140 Total Flow lb/hr 36667 36667 36667 36667 36667 36667 36667 5522 31144 31144 31144 1226 26022 3896 3896 31144 Temperature F 60 60 80 85 555 656 656 656 656 351 233 140 140 140 110 140 Pressure psia 15 85 84 3074 2984 2979 2969 2969 2969 1498 30 28 28 28 26 28 Vapor Frac 0 0 0 0 0 0 0 0 1 0 0 1 0 0 0 0 Mass Flow lb/hr H2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 CO2 0 0 0 0 0 0 2045 0 2045 2045 2045 1126 919 0 0 2045 H2O 27500 27500 27500 27500 27500 27500 27506 3301 24205 24205 24205 0 24049 156 156 24205 NH3 0 0 0 0 0 0 130 0 130 130 130 0 130 0 0 130 CH4 0 0 0 0 0 0 37 0 37 37 37 37 0 0 0 37 C2H6 0 0 0 0 0 0 12 0 12 12 12 12 0 0 0 12 C3H8 0 0 0 0 0 0 4 0 4 4 4 4 0 0 0 4 N-C4H10 0 0 0 0 0 0 34 0 34 34 34 34 0 0 0 34 N-PENTAN 0 0 0 0 0 0 12 0 12 12 12 12 0 0 0 12 METHANOL 0 0 0 0 0 0 218 0 218 218 218 0 218 0 0 218 ETHANOL 0 0 0 0 0 0 9 0 9 9 9 0 9 0 0 9 ACETONE 0 0 0 0 0 0 38 0 38 38 38 0 38 0 0 38 ACEACID 0 0 0 0 0 0 179 0 179 179 179 0 179 0 0 179 PROACID 0 0 0 0 0 0 65 0 65 65 65 0 65 0 0 65 ETHAMIN 0 0 0 0 0 0 96 0 96 96 96 0 96 0 0 96 2-PYRRLD 0 0 0 0 0 0 47 0 47 47 47 0 47 0 0 47 2-PIPERD 0 0 0 0 0 0 47 0 47 47 47 0 47 0 0 47 7-LACTAM 0 0 0 0 0 0 78 0 78 78 78 0 78 0 0 78 C5H9NS 0 0 0 0 0 0 245 0 245 245 245 0 89 156 156 245 TOLUENE 0 0 0 0 0 0 112 0 112 112 112 0 0 112 112 112 RYRO3ETM 0 0 0 0 0 0 198 0 198 198 198 0 0 198 198 198 PHENO4M 0 0 0 0 0 0 17 0 17 17 17 0 0 17 17 17 AMIPHENO 0 0 0 0 0 0 34 0 34 34 34 0 0 34 34 34 INDOLE 0 0 0 0 0 0 206 0 206 206 206 0 0 206 206 206 2-PYTENE 0 0 0 0 0 0 258 0 258 258 258 0 0 258 258 258 C15OLEF 0 0 0 0 0 0 258 0 258 258 258 0 0 258 258 258 MC12AMID 0 0 0 0 0 0 223 0 223 223 223 0 0 223 223 223 C16AMIDE 0 0 0 0 0 0 206 0 206 206 206 0 0 206 206 206 C18AMIDE 0 0 0 0 0 0 430 0 430 430 430 0 0 430 430 430 C16:0FA 0 0 0 0 0 0 304 0 304 304 304 0 12 292 292 304 C18:1FA 0 0 0 0 0 0 421 0 421 421 421 0 0 421 421 421 C13H18 0 0 0 0 0 0 344 0 344 344 344 0 0 344 344 344 HEVOIL1 0 0 0 0 0 0 418 0 418 418 418 0 47 371 371 418 HEVOIL2 0 0 0 0 0 0 215 0 215 215 215 0 0 215 215 215 SLUDGE 9167 9167 9167 9167 9167 9167 0 0 0 0 0 0 0 0 0 0 ASH 0 0 0 0 0 0 1377 1377 0 0 0 0 0 0 0 0 SOLID 0 0 0 0 0 0 845 845 0 0 0 0 0 0 0 0 Enthalpy mmBtu/hr -202.74 -202.73 -202.11 -201.70 -186.30 -181.06 -178.06 -23.28 -154.74 -170.14 -170.14 -4.45 -168.18 -3.35 -3.53 -175.76
A.4
671170
157
131
Fuel Gas From A100
156
A110 Combustor & Heat Transfer Circulation Oil
HTL Plant
Natural Gas
172
Exhaust
160
CombustorR-150
H-151Air
Air BlowerK-155
Exhaust BlowerK-150
Circ. Heat Transfer Oil
(Details not shown)
Heat from Hot OilTo A100 Feed Prep
Heat from Hot OilTo A100 HTL
171
A.5
131 156 157 160 170 171 172 Total Flow lb/hr 1226 374 11342 11342 9742 9742 9742 Temperature F 140 60 491 533 90 132 350 Pressure psia 28 450 15 17 15 18 16 Vapor Frac 1 1 1 1 1 1 1 Mass Flow lb/hr N2 0 2 7216 7216 7214 7214 7214 O2 0 0 368 368 2210 2210 2210 AR 0 0 123 123 123 123 123 CO2 1126 4 2433 2433 5 5 5 H2O 0 0 1202 1202 190 190 190 NH3 0 0 0 0 0 0 0 CH4 37 368 0 0 0 0 0 C2H6 12 0 0 0 0 0 0 C3H8 4 0 0 0 0 0 0 N-C4H10 34 0 0 0 0 0 0 N-PENTAN 12 0 0 0 0 0 0 HEXANE 0 0 0 0 0 0 0 METHANOL 0 0 0 0 0 0 0 ETHANOL 0 0 0 0 0 0 0 ACETONE 0 0 0 0 0 0 0 ACEACID 0 0 0 0 0 0 0 PROACID 0 0 0 0 0 0 0 ETHAMIN 0 0 0 0 0 0 0 2-PYRRLD 0 0 0 0 0 0 0 2-PIPERD 0 0 0 0 0 0 0 7-LACTAM 0 0 0 0 0 0 0 C5H9NS 0 0 0 0 0 0 0 TOLUENE 0 0 0 0 0 0 0 RYRO3ETM 0 0 0 0 0 0 0 PHENO4M 0 0 0 0 0 0 0 AMIPHENO 0 0 0 0 0 0 0 INDOLE 0 0 0 0 0 0 0 2-PYTENE 0 0 0 0 0 0 0 C15OLEF 0 0 0 0 0 0 0 MC12AMID 0 0 0 0 0 0 0 C16AMIDE 0 0 0 0 0 0 0 C18AMIDE 0 0 0 0 0 0 0 C16:0FA 0 0 0 0 0 0 0 C18:1FA 0 0 0 0 0 0 0 C13H18 0 0 0 0 0 0 0 HEVOIL1 0 0 0 0 0 0 0 HEVOIL2 0 0 0 0 0 0 0 Enthalpy, mmBtu/hr -4.45 -0.78 -15.04 -14.91 -1.09 -0.99 -0.46
A.6
HTL Aqueous From A100
Hydrated Lime
ClarifierCL-215
Lime Sludge
NH3 StripperT-220
THROX UnitTX-230
Water recycle to headworks
Air
Exhaust
Natural Gas
202
220
237
230
242
225
235
218
203
132
222 232
A200 HTL Aqueous TreatmentHTL Plant
H-230
K-230
A.7
132 202 203 218 220 222 225 230 232 235 237 242 Mass Flow lb/hr 25901 1314 175 25125 21139 58986 55000 55000 59161 59161 55000 2091 Temperature F 140 70 70 164 106 116 59 66 600 401 301 165 Pressure psia 28 15 25 60 15 15 15 15 15 15 15 60 Vapor Frac 0 0 1 0 0 1 1 1 1 1 1 0 Solid Frac 0 1 0 0 0 0 0 0 0 0 0 1 Volume Flow cuft/hr 2883 9 2476 415 341 885766 721495 706958 1658570 1346800 1023720 16 Mass Flow lb/hr H2O 23919 0 0 24097 20533 3911 347 347 4815 4815 347 268 N2 0 0 0 0 28 41249 41277 41277 41355 41355 41277 0 O2 0 0 0 0 22 12628 12650 12650 11285 11285 12650 0 AR 0 0 0 0 1 699 701 701 699 699 701 0 CO2 454 0 0 0 1 24 25 25 963 963 25 0 NH3 6 0 0 129 0 129 0 0 0 0 0 1 METHANE 0 0 175 0 0 0 0 0 0 0 0 0 METHANOL 218 0 0 216 10 206 0 0 0 0 0 2 ETHANOL 9 0 0 9 0 9 0 0 0 0 0 0 ETHAMIN 96 0 0 43 0 43 0 0 43 43 0 53 ACETONE 38 0 0 38 0 38 0 0 0 0 0 0 ACEACID 179 0 0 177 130 47 0 0 0 0 0 2 PROACID 65 0 0 64 62 2 0 0 0 0 0 1 2-PYRRLD 47 0 0 46 46 0 0 0 0 0 0 1 2-PIPERD 47 0 0 46 46 0 0 0 0 0 0 1 7-LACTAM 78 0 0 77 77 0 0 0 0 0 0 1 C5H9NS 89 0 0 88 88 0 0 0 0 0 0 1 C16:0FA 12 0 0 12 12 0 0 0 0 0 0 0 HEVOIL1 47 0 0 46 46 0 0 0 0 0 0 1 CA++ 0 0 0 7 7 0 0 0 0 0 0 0 K+ 7 0 0 7 7 0 0 0 0 0 0 0 NA+ 4 0 0 4 4 0 0 0 0 0 0 0 NH4+ 131 0 0 0 0 0 0 0 0 0 0 0 CL- 7 0 0 7 7 0 0 0 0 0 0 0 PO4--- 0 0 0 0 0 0 0 0 0 0 0 0 SO4-- 5 0 0 5 5 0 0 0 0 0 0 0 HCO3- 442 0 0 0 0 0 0 0 0 0 0 0 CO3-- 2 0 0 0 0 0 0 0 0 0 0 0 OH- 0 0 0 6 6 0 0 0 0 0 0 0 CA3PO4*2 0 0 0 0 0 0 0 0 0 0 0 0 CACO3 0 0 0 0 0 0 0 0 0 0 0 1759 CA(OH)2 0 1314 0 0 0 0 0 0 0 0 0 0
Enthalpy mmBtu/hr -168.01 -7.52 -0.35 -164.58 -140.62 -23.07 -2.35 -2.25 -23.42 -26.56 0.89 -10.96
A.8
HTL OilFrom Tanks
306
HydrotreaterR-310
307301
P-310
H-311
Makeup CompressorK-311
308
328
327
304303
Feed/Product HX-318A
Air Fin CondenserH-316
316
315
323
HP FlashV-311
LP FlashV-315
PSAV-300
318
Offgas to A350
332 320
To WWT
324
Lights ColumnT -310
334
Diesel ColumnT-330
338
337
339
To Gasoline Pool
To Diesel Pool
To A350 Hydrocracker
330
317
325
322
311
326
321E-310
Naphtha ColumnT -320
336
335
HydrogenFrom A400
Recycle CompressorK-310
HX-320
312
A310 HYDROTREATINGUpgrading Plant
319
A.9
301 304 306 307 308 311 312 315 316 317 320 321 322 324 325 326 327 328 330 331 332 334 335 336 337 338 339 341 Total Flow lb/hr 38961 43887 43887 38961 1988 43887 43887 43887 43887 43887 5171 36821 31476 31476 4128 2938 4926 4926 1988 36821 841 30635 6423 24212 12262 11950 5143 7066 Temperature F 80 335 591 92 354 750 581 246 300 110 78 110 78 220 110 110 209 244 140 78 108 409 205 560 487 656 108 110 Pressure psia 26 1525 1515 1530 710 1515 725 724 725 717 55 717 55 54 20 707 707 1530 336 55 48 50 25 25 19 19 20 717 Vapor Frac 0 1 1 0 1 1 1 1 1 1 0 0 0 0 1 1 1 1 1 0 1 0 0 0 0 0 1 1 Mass Flow lb/hr H2 0 4926 3284 0 1988 3284 3284 3284 3284 3284 0 19 1 1 326 2938 4926 4926 1988 19 1 0 0 0 0 0 346 3265 H2O 1560 1560 3525 1560 0 3525 3525 3525 3525 3525 3375 3383 6 6 142 0 0 0 0 3383 6 0 0 0 0 0 150 142 NH3 0 0 1969 0 0 1969 1969 1969 1969 1969 1259 1290 16 16 679 0 0 0 0 1290 16 0 0 0 0 0 711 679 CH4 0 0 850 0 0 850 850 850 850 850 0 31 6 6 819 0 0 0 0 31 6 0 0 0 0 0 850 819 C2H6 0 0 1028 0 0 1028 1028 1028 1028 1028 0 133 75 75 895 0 0 0 0 133 75 0 0 0 0 0 1027 895 C3H8 0 0 557 0 0 557 557 557 557 557 0 170 139 139 388 0 0 0 0 170 136 2 2 0 0 0 555 388 N-C4H10 0 0 188 0 0 188 188 188 188 188 1 104 99 99 84 0 0 0 0 104 57 42 42 0 0 0 145 84 N-PENTAN 0 0 1521 0 0 1521 1521 1521 1521 1521 6 1167 1147 1147 354 0 0 0 0 1167 287 860 860 0 0 0 655 354 HEXANE 0 0 354 0 0 354 354 354 354 354 0 317 315 315 37 0 0 0 0 317 39 276 276 0 0 0 78 37 C5H9NS 1559 1559 0 1559 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 TOLUENE 1117 1117 531 1117 0 531 531 531 531 531 0 516 516 516 15 0 0 0 0 516 23 492 492 0 0 0 39 15 RYRO3ETM 1976 1976 0 1976 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 PHENO4M 172 172 0 172 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 AMIPHENO 344 344 0 344 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 INDOLE 2062 2062 0 2062 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 2-PYTENE 2578 2578 0 2578 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C15OLEF 2578 2578 0 2578 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 MC12AMID 2234 2234 0 2234 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C16AMIDE 2062 2062 0 2062 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C18AMIDE 4296 4296 0 4296 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C16:0FA 2921 2921 496 2921 0 496 496 496 496 496 0 496 495 495 0 0 0 0 0 496 0 495 0 495 25 470 0 0 C18:1FA 4210 4210 0 4210 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C13H18 3437 3437 0 3437 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 HEVOIL1 3708 3708 0 3708 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 HEVOIL2 2148 2148 0 2148 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 2MPENTA 0 0 460 0 0 460 460 460 460 460 1 397 394 394 63 0 0 0 0 397 61 334 334 0 0 0 126 63 CC5-METH 0 0 248 0 0 248 248 248 248 248 1 225 224 224 22 0 0 0 0 225 24 200 200 0 0 0 47 22 2MHEXAN 0 0 460 0 0 460 460 460 460 460 1 433 431 431 28 0 0 0 0 433 33 398 398 0 0 0 61 28 CC6-METH 0 0 177 0 0 177 177 177 177 177 0 170 169 169 7 0 0 0 0 170 9 160 160 0 0 0 17 7 PIPERDIN 0 0 71 0 0 71 71 71 71 71 16 70 54 54 0 0 0 0 0 70 2 52 52 0 0 0 2 0 3MHEPTA 0 0 354 0 0 354 354 354 354 354 0 345 345 345 8 0 0 0 0 345 12 333 333 0 0 0 21 8 OCTANE 0 0 248 0 0 248 248 248 248 248 0 243 243 243 4 0 0 0 0 243 7 237 237 0 0 0 11 4 ETHCYC6 0 0 265 0 0 265 265 265 265 265 0 261 261 261 4 0 0 0 0 261 6 255 255 0 0 0 10 4 ETHBENZ 0 0 265 0 0 265 265 265 265 265 0 262 262 262 3 0 0 0 0 262 6 256 256 0 0 0 9 3 O-XYLENE 0 0 425 0 0 425 425 425 425 425 0 421 420 420 4 0 0 0 0 421 7 413 413 0 0 0 11 4 C9H20 0 0 602 0 0 602 602 602 602 602 0 597 597 597 4 0 0 0 0 597 7 590 590 0 0 0 12 4 CC6-PRO 0 0 460 0 0 460 460 460 460 460 0 457 457 457 3 0 0 0 0 457 5 452 451 0 0 0 8 3 C3BENZ 0 0 283 0 0 283 283 283 283 283 0 282 281 281 2 0 0 0 0 282 3 278 278 0 0 0 5 2 4MNONAN 0 0 283 0 0 283 283 283 283 283 0 282 282 282 1 0 0 0 0 282 2 280 279 1 1 0 4 1 C10H22 0 0 354 0 0 354 354 354 354 354 0 353 353 353 1 0 0 0 0 353 2 351 337 14 14 0 3 1 C4BENZ 0 0 531 0 0 531 531 531 531 531 0 530 530 530 1 0 0 0 0 530 2 527 175 352 352 0 4 1 C10H12 0 0 602 0 0 602 602 602 602 602 0 601 600 600 1 0 0 0 0 601 2 598 2 596 596 0 3 1 C12H26 0 0 425 0 0 425 425 425 425 425 0 425 425 425 0 0 0 0 0 425 0 424 0 424 424 0 1 0 1234NA 0 0 708 0 0 708 708 708 708 708 0 707 707 707 1 0 0 0 0 707 2 705 0 705 705 0 2 1 CC6-HEX 0 0 708 0 0 708 708 708 708 708 0 708 708 708 0 0 0 0 0 708 1 707 0 707 707 0 1 0 14DMNAPH 0 0 885 0 0 885 885 885 885 885 0 885 884 884 0 0 0 0 0 885 1 884 0 884 884 0 1 0 C7BENZ 0 0 1062 0 0 1062 1062 1062 1062 1062 0 1062 1061 1061 0 0 0 0 0 1062 0 1061 0 1061 1061 0 1 0 C8BENZ 0 0 814 0 0 814 814 814 814 814 0 814 814 814 0 0 0 0 0 814 0 814 0 814 814 0 0 0 TETMC16 0 0 2124 0 0 2124 2124 2124 2124 2124 0 2124 2124 2124 0 0 0 0 0 2124 0 2124 0 2124 675 1448 0 0 C22H42O4 0 0 708 0 0 708 708 708 708 708 0 708 708 708 0 0 0 0 0 708 0 708 0 708 0 708 0 0 C15H32 0 0 708 0 0 708 708 708 708 708 0 708 708 708 0 0 0 0 0 708 0 708 0 708 707 1 0 0 C16H34 0 0 2124 0 0 2124 2124 2124 2124 2124 0 2124 2124 2124 0 0 0 0 0 2124 0 2124 0 2124 2053 71 0 0 C17H36 0 0 3540 0 0 3540 3540 3540 3540 3540 0 3539 3539 3539 0 0 0 0 0 3539 0 3539 0 3539 2247 1292 0 0 C18H38 0 0 3186 0 0 3186 3186 3186 3186 3186 0 3186 3186 3186 0 0 0 0 0 3186 0 3186 0 3186 930 2255 0 0 C20H42 0 0 2124 0 0 2124 2124 2124 2124 2124 0 2124 2124 2124 0 0 0 0 0 2124 0 2124 0 2124 65 2059 0 0 C22H46 0 0 531 0 0 531 531 531 531 531 0 531 531 531 0 0 0 0 0 531 0 531 0 531 1 530 0 0 C24H50 0 0 1239 0 0 1239 1239 1239 1239 1239 0 1239 1239 1239 0 0 0 0 0 1239 0 1239 0 1239 0 1239 0 0 C28H58 0 0 885 0 0 885 885 885 885 885 0 885 885 885 0 0 0 0 0 885 0 885 0 885 0 885 0 0 C30H62 0 0 283 0 0 283 283 283 283 283 0 283 283 283 0 0 0 0 0 283 0 283 0 283 0 283 0 0 C32H66 0 0 708 0 0 708 708 708 708 708 0 708 708 708 0 0 0 0 0 708 0 708 0 708 0 708 0 0 NH4HS 0 0 689 0 0 689 689 689 689 689 459 459 1 1 229 0 0 0 0 459 0 1 0 1 0 1 229 229 2*NH4CO3 0 0 49 0 0 49 49 49 49 49 49 49 0 0 0 0 0 0 0 49 0 0 0 0 0 0 0 0 Enthalpy, mmBtu/hr -35.88 -24.19 -31.70 -35.40 1.93 -24.48 -31.73 -46.31 -44.00 -52.51 -26.52 -49.47 -23.81 -21.50 -5.46 0.36 2.29 2.97 0.44 -49.47 -0.75 -17.05 -4.12 -11.23 -4.72 -6.28 -6.40 -5.13
A.10
Heavy OilFrom A310
HydrocrackerR-350
HydrogenFrom A400
P-350
H-351Make-up Compressor
K-350
V-380
Naphtha SplitterT-350
385
To Gasoline Pool
To Diesel Pool
386
Trim HeaterH-352
365
355
383
389
377
362
381
353
350
351
380
361
Off-Gas From A310
Air Fin CondenserH-371
T-350 ReboilerHX-381
V-375
352
376
378
Make-up CompressorK-355
339
370
374 372
368
To Ammonia Scrubber
366
A350 HYDROCRACKINGUpgrading Plant
A.11
339 350 351 352 353 355 361 362 365 366 368 370 372 374 376 377 378 380 381 383 385 386 389 Total Flow lb/hr 5143 11950 125 125 12676 12676 11950 12676 12676 5557 182 12676 12676 12676 752 602 150 602 11924 11742 82 716 10944 Temperature F 108 653 140 297 584 700 661 741 753 135 137 655 140 300 140 140 140 148 140 137 140 140 520 Pressure psia 20 19 336 1040 1035 1030 1035 1010 1010 20 30 1009 1006 1008 1006 1006 1006 1040 1006 30 20 20 24 Vapor Frac 1 0 1 1 1 1 0 1 1 1 1 1 1 1 1 1 1 1 0 0 1 0 0 Mass Flow lb/hr H2 346 0 125 125 420 420 0 420 377 428 8 377 377 377 368 295 74 295 8 0 0 0 0 CO2 0 0 0 0 262 262 0 262 430 168 91 430 430 430 327 262 65 262 103 12 11 0 0 H2O 150 0 0 0 12 12 0 12 78 216 51 78 78 78 15 12 3 12 63 12 11 1 0 NH3 711 0 0 0 0 0 0 0 0 711 0 0 0 0 0 0 0 0 0 0 0 0 0 CH4 850 0 0 0 0 0 0 0 0 850 0 0 0 0 0 0 0 0 0 0 0 0 0 C2H6 1027 0 0 0 0 0 0 0 0 1027 0 0 0 0 0 0 0 0 0 0 0 0 0 C3H8 555 0 0 0 0 0 0 0 0 555 0 0 0 0 0 0 0 0 0 0 0 0 0 N-C4H10 145 0 0 0 26 26 0 26 137 215 23 137 137 137 33 26 7 26 104 82 40 41 1 N-PENTAN 655 0 0 0 0 0 0 0 0 655 0 0 0 0 0 0 0 0 0 0 0 0 0 HEXANE 78 0 0 0 1 1 0 1 15 80 1 15 15 15 1 1 0 1 15 14 2 12 0 TOLUENE 39 0 0 0 0 0 0 0 0 39 0 0 0 0 0 0 0 0 0 0 0 0 0 C16:0FA 0 470 0 0 470 470 470 470 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 2MPENTA 126 0 0 0 0 0 0 0 0 126 0 0 0 0 0 0 0 0 0 0 0 0 0 CC5-METH 47 0 0 0 0 0 0 0 0 47 0 0 0 0 0 0 0 0 0 0 0 0 0 2MHEXAN 61 0 0 0 0 0 0 0 0 61 0 0 0 0 0 0 0 0 0 0 0 0 0 HEPTANE 0 0 0 0 1 1 0 1 75 5 1 75 75 75 1 1 0 1 74 72 4 63 6 CC6-METH 17 0 0 0 0 0 0 0 0 17 0 0 0 0 0 0 0 0 0 0 0 0 0 PIPERDIN 2 0 0 0 0 0 0 0 0 2 0 0 0 0 0 0 0 0 0 0 0 0 0 3MHEPTA 21 0 0 0 3 3 0 3 440 35 4 440 440 440 4 3 1 3 436 432 10 344 78 OCTANE 11 0 0 0 1 1 0 1 166 15 1 166 166 166 1 1 0 1 165 163 3 123 38 ETHCYC6 10 0 0 0 0 0 0 0 0 10 0 0 0 0 0 0 0 0 0 0 0 0 0 ETHBENZ 9 0 0 0 0 0 0 0 0 9 0 0 0 0 0 0 0 0 0 0 0 0 0 O-XYLENE 11 0 0 0 0 0 0 0 0 11 0 0 0 0 0 0 0 0 0 0 0 0 0 C9H20 12 0 0 0 1 1 0 1 253 14 1 253 253 253 1 1 0 1 252 252 1 114 137 CC6-PRO 8 0 0 0 0 0 0 0 0 8 0 0 0 0 0 0 0 0 0 0 0 0 0 C3BENZ 5 0 0 0 0 0 0 0 0 5 0 0 0 0 0 0 0 0 0 0 0 0 0 4MNONAN 4 0 0 0 0 0 0 0 0 4 0 0 0 0 0 0 0 0 0 0 0 0 0 C10H22 3 0 0 0 0 0 0 0 391 4 0 391 391 391 0 0 0 0 390 390 0 19 370 C4BENZ 4 0 0 0 0 0 0 0 0 4 0 0 0 0 0 0 0 0 0 0 0 0 0 C11H24 0 0 0 0 0 0 0 0 501 0 0 501 501 501 0 0 0 0 501 501 0 0 501 C10H12 3 0 0 0 0 0 0 0 0 3 0 0 0 0 0 0 0 0 0 0 0 0 0 C12H26 1 0 0 0 0 0 0 0 563 1 0 563 563 563 0 0 0 0 563 563 0 0 563 1234NA 2 0 0 0 0 0 0 0 0 2 0 0 0 0 0 0 0 0 0 0 0 0 0 CC6-HEX 1 0 0 0 0 0 0 0 0 1 0 0 0 0 0 0 0 0 0 0 0 0 0 14DMNAPH 1 0 0 0 0 0 0 0 0 1 0 0 0 0 0 0 0 0 0 0 0 0 0 C13H28 0 0 0 0 0 0 0 0 520 0 0 520 520 520 0 0 0 0 520 520 0 0 520 C7BENZ 1 0 0 0 0 0 0 0 0 1 0 0 0 0 0 0 0 0 0 0 0 0 0 TETMC16 0 1448 0 0 1448 1448 1448 1448 1448 0 0 1448 1448 1448 0 0 0 0 1448 1448 0 0 1448 C22H42O4 0 708 0 0 708 708 708 708 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C14H30 0 0 0 0 0 0 0 0 384 0 0 384 384 384 0 0 0 0 384 384 0 0 384 C15H32 0 1 0 0 1 1 1 1 337 0 0 337 337 337 0 0 0 0 337 337 0 0 337 C16H34 0 71 0 0 71 71 71 71 695 0 0 695 695 695 0 0 0 0 695 695 0 0 695 C17H36 0 1292 0 0 1292 1292 1292 1292 1423 0 0 1423 1423 1423 0 0 0 0 1423 1423 0 0 1423 C18H38 0 2255 0 0 2255 2255 2255 2255 2324 0 0 2324 2324 2324 0 0 0 0 2324 2324 0 0 2324 C19H40 0 0 0 0 0 0 0 0 35 0 0 35 35 35 0 0 0 0 35 35 0 0 35 C20H42 0 2059 0 0 2059 2059 2059 2059 2074 0 0 2074 2074 2074 0 0 0 0 2074 2074 0 0 2074 C21H44 0 0 0 0 0 0 0 0 9 0 0 9 9 9 0 0 0 0 9 9 0 0 9 C22H46 0 530 0 0 530 530 530 530 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C24H50 0 1239 0 0 1239 1239 1239 1239 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C28H58 0 885 0 0 885 885 885 885 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C30H62 0 283 0 0 283 283 283 283 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C32H66 0 708 0 0 708 708 708 708 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 NH4HS 229 1 0 0 1 1 1 1 1 229 0 1 1 1 0 0 0 0 0 0 0 0 0 Enthalpy, mmBtu/hr -6.40 -6.36 0.03 0.10 -7.20 -5.84 -6.27 -5.30 -5.30 -7.49 -0.67 -6.67 -11.85 -10.49 -1.29 -1.03 -0.26 -1.02 -10.56 -9.93 -0.16 -0.66 -6.38
A.12
C389 391
392
A350 Naphtha Naphtha Storage
386
A310 Naphtha
337
Diesel Storage
393
A310 Diesel
390
Air CoolerH-365
A350 Diesel
389
Air CoolerH-360
Heat RecoveryHX-385
335
A300 PRODUCT COOLINGUpgrading Plant
392
391 394
A.13
335 337 386 389 390 391 392 393 394 Total Flow lb/hr 6423 12262 716 10944 7140 23206 7140 23206 23206 Temperature F 205 487 140 520 192 496 140 140 300 Pressure psia 25 19 20 24 20 19 19 17 18 Vapor Frac 0 0 0 0 0 0 0 0 0 Mass Flow lb/hr H2O 0 0 1 0 1 0 1 0 0 C3H8 2 0 0 0 2 0 2 0 0 N-C4H10 42 0 41 1 83 1 83 1 1 N-PENTAN 860 0 0 0 860 0 860 0 0 HEXANE 276 0 12 0 288 0 288 0 0 TOLUENE 492 0 0 0 492 0 492 0 0 C16:0FA 0 25 0 0 0 25 0 25 25 2MPENTA 334 0 0 0 334 0 334 0 0 CC5-METH 200 0 0 0 200 0 200 0 0 2MHEXAN 398 0 0 0 398 0 398 0 0 HEPTANE 0 0 63 6 63 6 63 6 6 CC6-METH 160 0 0 0 160 0 160 0 0 PIPERDIN 52 0 0 0 52 0 52 0 0 3MHEPTA 333 0 344 78 677 78 677 78 78 OCTANE 237 0 123 38 360 38 360 38 38 ETHCYC6 255 0 0 0 255 0 255 0 0 ETHBENZ 256 0 0 0 256 0 256 0 0 O-XYLENE 413 0 0 0 413 0 413 0 0 C9H20 590 0 114 137 703 137 703 137 137 CC6-PRO 451 0 0 0 451 0 451 0 0 C3BENZ 278 0 0 0 278 0 278 0 0 4MNONAN 279 1 0 0 279 1 279 1 1 C10H22 337 14 19 370 356 384 356 384 384 C4BENZ 175 352 0 0 175 352 175 352 352 C11H24 0 0 0 501 0 501 0 501 501 C10H12 2 596 0 0 2 596 2 596 596 C12H26 0 424 0 563 0 987 0 987 987 1234NA 0 705 0 0 0 705 0 705 705 CC6-HEX 0 707 0 0 0 707 0 707 707 14DMNAPH 0 884 0 0 0 884 0 884 884 C13H28 0 0 0 520 0 520 0 520 520 C7BENZ 0 1061 0 0 0 1061 0 1061 1061 C8BENZ 0 814 0 0 0 814 0 814 814 TETMC16 0 675 0 1448 0 2124 0 2124 2124 C14H30 0 0 0 384 0 384 0 384 384 C15H32 0 707 0 337 0 1044 0 1044 1044 C16H34 0 2053 0 695 0 2748 0 2748 2748 C17H36 0 2247 0 1423 0 3670 0 3670 3670 C18H38 0 930 0 2324 0 3255 0 3255 3255 C19H40 0 0 0 35 0 35 0 35 35 C20H42 0 65 0 2074 0 2139 0 2139 2139 C21H44 0 0 0 9 0 9 0 9 9 C22H46 0 1 0 0 0 1 0 1 1 Enthalpy, mmBtu/hr -4.12 -4.72 -0.66 -6.38 -4.78 -11.10 -5.01 -16.36 -14.27
A.14
NH3 Scrubber
Offgas From A350
366
To A400
367
Water
To WWT
A300 AMMONIA SCRUBBERUpgrading Plant
A.15
366 WATER 367 To WWT
Total Flow lb/hr 5557 15313 4905 15965
Temperature F 135 90 114 116
Pressure psia 20 30 20 20
Vapor Frac 1 0 1 0
Mass Flow lb/hr
H2 428 0 428 0
CO2 168 0 167 0
H2O 216 15313 472 15057
NH3 711 0 36 674
CH4 850 0 850 0
C2H6 1027 0 1027 0
C3H8 555 0 555 0
N-C4H10 215 0 215 0
N-PENTAN 655 0 654 1
HEXANE 80 0 80 0
TOLUENE 39 0 38 0
C16:0FA 0 0 0 0
2MPENTA 126 0 125 0
CC5-METH 47 0 47 0
2MHEXAN 61 0 61 0
HEPTANE 5 0 5 0
CC6-METH 17 0 17 0
PIPERDIN 2 0 1 1
3MHEPTA 35 0 35 0
OCTANE 15 0 15 0
ETHCYC6 10 0 10 0
ETHBENZ 9 0 9 0
O-XYLENE 11 0 11 0
C9H20 14 0 14 0
CC6-PRO 8 0 8 0
C3BENZ 5 0 5 0
4MNONAN 4 0 4 0
C10H22 4 0 4 0
C4BENZ 4 0 4 0
C11H24 0 0 0 0
C10H12 3 0 3 0
C12H26 1 0 1 0
1234NA 2 0 2 0
CC6-HEX 1 0 1 0
14DMNAPH 1 0 1 0
C7BENZ 1 0 1 0
NH4HS 229 0 0 229
Enthalpy, mmBtu/hr -7.49 -104.21 -7.88 -103.82
A.16
426
401
428
427
416
419
421
330
Steam Reformer
420
Air
367A350 Offgas from
NH3 Scrubber
Flue Gas
Condensate to BFW makeup
411
460
468
402
410
413
439
423
HydrogenTo A350
HydrogenTo A310
351
434417
Natural Gas
Steam Drum
Pre-Reformer
Sulfur Removal
Water Gas Shift
PSA
425
415
A400 HYDROGEN PLANTUpgrading Plant
461
Condensate to WWT
436
H-415
H-414
K-400
K-462
K-461
H-420
H-425
H-430
A.17
330 351 367 401 402 410 411 413 415 417 419 420 421 423 425 426 427 428 434 436 439 460 461 468 Total Flow lb/hr 1988 125 4905 1203 1475 27575 27575 27575 21563 27575 27575 27575 15632 27575 2113 11943 13519 58212 4537 368 27575 58212 58212 72934 Temperature F 140 140 114 60 60 542 900 957 700 300 1562 140 140 572 140 140 140 400 291 114 1050 90 175 322 Pressure psia 336 336 20 450 450 450 450 410 659 348 359 346 346 359 336 346 20 21 450 20 409 15 21 16 Vapor Frac 1 1 1 1 1 1 1 1 1 1 1 1 1 1 1 0 1 1 1 0 1 1 1 1 Mass Flow lb/hr H2 1988 125 428 0 0 428 428 434 0 2347 2066 2347 2347 2066 2113 0 235 0 428 0 434 0 0 0 CO 0 0 0 0 0 0 0 41 0 593 4505 593 593 4505 0 0 593 0 0 0 41 0 0 0 N2 0 0 0 7 8 8 8 35 0 36 36 36 36 36 0 0 36 43105 0 0 35 43105 43105 43147 O2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 13207 0 0 0 13207 13207 2199 AR 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 735 0 0 0 735 735 735 CO2 0 0 167 13 16 184 184 1638 0 11427 5279 11427 11418 5279 0 9 11418 29 167 0 1638 29 29 18448 H2O 0 0 472 0 0 21731 21731 20514 21563 12145 14662 12145 211 14662 0 11934 211 1136 168 303 20514 1136 1136 8403 H2S 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 SO2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 NH3 0 0 36 0 0 36 36 4 0 2 2 2 2 2 0 0 2 0 36 0 4 0 0 2 CH4 0 0 850 1183 1450 2300 2300 4909 0 1024 1024 1024 1024 1024 0 0 1024 0 849 0 4909 0 0 0 C2H6 0 0 1027 0 0 1027 1027 0 0 0 0 0 0 0 0 0 0 0 1027 1 0 0 0 0 C3H8 0 0 555 0 0 554 554 0 0 0 0 0 0 0 0 0 0 0 554 1 0 0 0 0 N-C4H10 0 0 215 0 0 214 214 0 0 0 0 0 0 0 0 0 0 0 214 1 0 0 0 0 N-PENTAN 0 0 654 0 0 646 646 0 0 0 0 0 0 0 0 0 0 0 646 8 0 0 0 0 HEXANE 0 0 80 0 0 78 78 0 0 0 0 0 0 0 0 0 0 0 78 2 0 0 0 0 TOLUENE 0 0 38 0 0 34 34 0 0 0 0 0 0 0 0 0 0 0 34 4 0 0 0 0 C16:0FA 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 2MPENTA 0 0 125 0 0 123 123 0 0 0 0 0 0 0 0 0 0 0 123 3 0 0 0 0 CC5-METH 0 0 47 0 0 46 46 0 0 0 0 0 0 0 0 0 0 0 46 2 0 0 0 0 2MHEXAN 0 0 61 0 0 58 58 0 0 0 0 0 0 0 0 0 0 0 58 3 0 0 0 0 HEPTANE 0 0 5 0 0 5 5 0 0 0 0 0 0 0 0 0 0 0 5 0 0 0 0 0 CC6-METH 0 0 17 0 0 15 15 0 0 0 0 0 0 0 0 0 0 0 15 1 0 0 0 0 3MHEPTA 0 0 35 0 0 30 30 0 0 0 0 0 0 0 0 0 0 0 30 5 0 0 0 0 OCTANE 0 0 15 0 0 13 13 0 0 0 0 0 0 0 0 0 0 0 13 3 0 0 0 0 ETHCYC6 0 0 10 0 0 8 8 0 0 0 0 0 0 0 0 0 0 0 8 2 0 0 0 0 ETHBENZ 0 0 9 0 0 7 7 0 0 0 0 0 0 0 0 0 0 0 7 2 0 0 0 0 O-XYLENE 0 0 11 0 0 8 8 0 0 0 0 0 0 0 0 0 0 0 8 3 0 0 0 0 C9H20 0 0 14 0 0 9 9 0 0 0 0 0 0 0 0 0 0 0 9 5 0 0 0 0 CC6-PRO 0 0 8 0 0 5 5 0 0 0 0 0 0 0 0 0 0 0 5 3 0 0 0 0 C3BENZ 0 0 5 0 0 3 3 0 0 0 0 0 0 0 0 0 0 0 3 2 0 0 0 0 4MNONAN 0 0 4 0 0 2 2 0 0 0 0 0 0 0 0 0 0 0 2 2 0 0 0 0 C10H22 0 0 4 0 0 2 2 0 0 0 0 0 0 0 0 0 0 0 2 2 0 0 0 0 C4BENZ 0 0 4 0 0 1 1 0 0 0 0 0 0 0 0 0 0 0 1 2 0 0 0 0 C10H12 0 0 3 0 0 1 1 0 0 0 0 0 0 0 0 0 0 0 1 2 0 0 0 0 C12H26 0 0 1 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 1 0 0 0 0 1234NA 0 0 2 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 2 0 0 0 0 CC6-HEX 0 0 1 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 1 0 0 0 0 14DMNAPH 0 0 1 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 1 0 0 0 0 C13H28 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C7BENZ 0 0 1 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 1 0 0 0 0 TETMC16 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C22H42O4 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C15H32 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C16H34 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C17H36 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C18H38 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C19H40 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C20H42 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C21H44 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C22H46 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C24H50 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C28H58 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C30H62 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 C32H66 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 NH4HS 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 0 Enthalpy, mmBtu/hr 0.44 0.03 -7.88 -2.44 -2.99 -127.43 -121.36 -121.36 -118.93 -118.61 -87.74 -128.91 -47.49 -106.37 0.47 -81.42 -47.92 -2.04 -5.50 -2.11 -119.71 -6.49 -5.27 -114.72
A.18
618
426
709
648
Boiler Feedwater Makeup
Steam TurbineE-601
Steam System A600
645
600
647
Blowdown
Steam to A400
CondenserHX-600
A300, A400 Process Heat
415
Cooling Water Service – A700
Cooling Tower Makeup
A310, 350 Service Return
730
710Blowdown
Losses
Cooling Water to A310,350
715
713
760
A600 Service Return
A400 Condensate
A600 STEAM SYSTEMA700 COOLING WATER
Upgrading Plant
633
709
618
A.19
415 426 600 618 633 645 648 647 709 710 713 715 730 760
Total Flow lb/hr 21563 11943 851 10479 42546 20124 20124 20124 1016260 23485 4453 1036140 100370 935773
Temperature F 700 140 112 60 232 700 148 148 90 60 89 89 110 110
Pressure psia 659 346 15 15 22 659 4 4 15 15 15 75 60 60
Vapor Frac 1 0 0 0 0 1 0 1 0 0 0 0 0 0
Mass Flow lb/hr
H2 0 0 0 0 0 0 0 0 0 0 0 0 0
CO 0 0 0 0 0 0 0 0 0 0 0 0 0
N2 0 0 0 0 0 0 0 0 0 0 0 0 0
CO2 9 0 0 9 0 0 0 0 0 0 0 0 0
H2O 21563 11934 851 10479 42538 20124 20124 20124 1016260 23485 4453 1036140 100370 935773
NH3 0 0 0 0 0 0 0 0 0 0 0 0 0
CH4 0 0 0 0 0 0 0 0 0 0 0 0 0
C2H6 0 0 0 0 0 0 0 0 0 0 0 0 0
Enthalpy, mmBtu/hr -118.93 -81.42 -5.78 -71.66 -283.58 -111.05 -135.85 -117.56 -6990.40 -162.35 -30.63 -7127.60 -688.07 -6415.00
A.20
Heat Exchanger Duty Summary Power Summary
Power Summary
Process Flow Diagram Equip ID Duty, mmBTU/hr Equip ID HP A100 - HTL
H-100 0.61 P101a 3.34
Hx-100 15.41 P101b 160.87 H-105 5.24 R-101 3 HX-102 5.26 H-110 0.18 A110 - Combustor
H-151 0.52 K-150 50.92
A200 - Aqueous Treatment K-155 39.09
H-230 3.14 K-230 36.01
A310 - Hydrotreating H-311 7.25 P-310 106.57 HX-318A 2.31 K-311 585.59 H-316 4.84 K-310 1497.52 E-310 1.36 T-310 Condenser 0.64 T-310 Reboiler 4.34 T-320 Condenser 4.55 T-320 Reboiler 6.26 T-336 Condenser 2.64 T-336 Reboiler 2.87 A350 – Hydrocracking
H-351 1.36 P-350 35.53
H-352 0.53 K-350 53.91 H-371 1.36 K-355 3.51 T-350 Condenser 0.24 T-350 Reboiler 2.98 A300 - Product Cooling H-360 0.23 HX-385 3.18 H-365 2.08 A400 - Hydrogen Plant H-414 6.07 K-400 841.39 H-415 1.64 K-461 477.53 H-420 18.63 K-462 277.69 H-425 12.24 H-430 10.3 A600 - Steam System HX-600 18.29 E-601 2556.50
B.1
- Equipment Cost Details
HTL Plant:
Equipment
Number
Number
Required
Number
Spares Equipment Name
Scaling
Stream
Original
Equipment
Stream Flow
New
Flows
stream
flow units
Size
Ratio
Original Equip
Cost (per unit)
Base
Year
COST BASIS:
installed (i) or
bare (b)
Total Original
Equip Cost
(Req'd &
Spare) in
Base Year
Scaling
Exponent
Scaled Cost
in Base Year
Installation
Factor
Installed Cost
in Base Year
Installed Cost
in 2014$
Scaled
Uninstalled
Cost in
2014$ Ref
A 10 Feed Prep - Sludge Dewatering and Grinding
C-101
0 Bowl Centrifuge - GEA
raw sludge
rate 26,417 29,296 gal/hr 1.11 $700,000 2016 b $700,000 0.6 744,817 1.8 1,340,670 1,425,808 $792,115 2
A10 Total 1,425,808 792,115
A100 HTL Oil ProductionP-101 1 0 Twin Screw Feeder and Feed Pump S101 575 69 gal/min 0.12 $1,094,100 2011 b $1,094,100 0.8 200,098 2.3 460,226 452,683 $196,819 1HX-100 1 0 Feed/Product Exchangers Area 35,540 1,284 ft2 0.04 $31,860,000 2012 b $31,860,000 0.7 3,117,016 2.2 6,857,436 6,757,730 $3,071,695 1H-105 1 0 Feed Trim Heater duty 32.1 5.2 mmBtu/hr 0.16 $998,850 2012 b $998,850 0.7 279,350 2.2 614,571 605,635 $275,289 1R-101 1 0 HTL Reactor S101 33,333 36,667 lb/h 1.10 $770,992 1Q 2014 b $770,992 1 848,109 1.2 975,326 980,431 $852,549 3K-110 1 0 Reactor K/O Drum S101 304,939 36,667 lb/h 0.12 $5,700,243 2011 b $5,700,243 0.7 1,294,010 2.0 2,588,021 2,545,601 $1,272,801 1
1 0 Hot Oil System duty 64 8.6 mmBtu/hr 0.14 $1,265,600 2011 b $1,265,600 0.6 380,997 1.4 545,968 537,020 $374,752 1
1 0 Dowtherm
capacity of
system 23,200 3,137 gal 0.14 $782,847 2011 b $782,847 1 105,857 1.0 105,857 104,122 $104,122 1HX-102 1 0 Reactor Product Air Cooler duty 5.12 5.26 mmBtu/hr 1.03 $63,200 1Q 2014 b $63,200 0.7 64,405 2.0 127,521 128,189 $64,742 3F-105 1 0 Solids Filter S101 2,420 69 gal/min 0.03 $1,017,998 2011 b $1,017,998 0.6 120,198 1.7 204,337 200,988 $118,228 1V-110 1 0 Oil/Water Separator S101 2,420 69 gal/min 0.03 $2,653,959 2011 b $2,653,959 0.7 219,485 2.0 438,969 431,774 $215,887 1
H-110 1 0 Biocrude Cooler duty 0.170 0.181 mmBtu/hr 1.06 $63,200 1Q 2014 b $63,200 0.7 66,008 5.4 356,442 358,308 $66,353 3
A100 Total 13,102,481 6,613,237
A200 HTL Water TreatmentCL-215 1 0 Lime Softening S101 34,000 36,667 lb/hr 1.08 $410,100 1Q2014 b $410,100 0.65 430,730 2.4 1,033,752 1,039,163 $432,985 5T-220 1 0 Ammonia Stripping S101 34,000 36,667 lb/hr 1.08 $596,800 1Q2014 b $596,800 0.65 626,822 2.4 1,504,373 1,512,248 $630,103 5
TX-230 1 0 THROX Unit S101 34,000 36,667 lb/hr 1.08 $210,900 1Q2014 b $210,900 0.65 221,509 2.4 531,622 534,405 $222,669 5
A200 $542,000 $3,085,816 $1,285,757
A700 OSBL - including cooling water system
1 0 Cooling Tower System (packaged) circ rate 150 18 gal/min 0.12 $5,900 1Q 2014 b $5,900 0.6 1,659 10.1 16,756 16,844 $1,668 31 1 Cooling Water Pump circ rate 16 18 gal/min 1.13 $4,400 1Q 2014 b $8,800 0.6 9,477 4 37,907 38,105 $9,526 31 0 Plant Air Compressor S101 - dry 2,205 110 ton/day 0.05 $87,922 2007 b $87,922 0.3 35,770 1.57 56,158 61,577 $39,221 41 0 Firewater Pump S101 - dry 2,205 110 ton/day 0.05 $23,043 2007 b $23,043 0.3 9,375 3.7 34,686 38,033 $10,279 4
1 0 Instrument Air Dryer S101 - dry 2,205 110 ton/day 0.05 $8,349 2002 b $8,349 0.6 1,382 2.47 3,413 4,971 $2,012 4
1 0 Plant Air Receiver S101 - dry 2,205 110 ton/day 0.05 $21,005 2007 b $21,005 0.65 2,993 5.44 16,280 17,851 $3,281 4
1 0 Firewater Storage Tank S101 - dry 2,205 110 ton/day 0.05 $229,900 2007 b $229,900 0.65 32,755 1.46 47,822 52,437 $35,916 4
1 Biocrude Storage - 3 day S136 31,719 34,419 gal 1.09 $280,900 1Q 2014 b $280,900 0.65 296,216 1.3 385,081 387,097 $297,767 3
A700 Total Subtotal $616,915 $399,671
References Total Equipment Cost $18,231,020 $9,090,780
1 Knorr et al. 2013
2 Vendor Budget Estimate
3 Aspen Capital Cost Estimator, 8.8
4 Dutta et al. 2011
5 Aspen Process Economic Analyzer, V.8.8
B.2
Upgrading Plant:
Equipment
Number
Number
Required
Number
Spares Equipment Name Scaling Stream
Original
Equipment
Stream Flow
New
Flows
stream
flow units
Size
Ratio
Original
Equip Cost
(per unit)
Base
Year
COST
BASIS:
installed (i)
or bare (b)
Total Original
Equip Cost
(Req'd & Spare)
in Base Year
Scaling
Exponent
Scaled Cost
in Base Year
Installation
Factor
Installed Cost
in Base Year
Installed
Cost in
2014$
Scaled
Uninstalled
Cost in 2014$ Ref
A300 HTL Oil Upgrading and Product Separation
A310 HTL Oil Hydrotreating & Separations1 0 Desalter S301 2,732 2,732 bpd feed 1.00 $768,541 2007 i $768,541 0.75 768,541 2.5 768,541 842,704 $341,176 71 0 Hydrotreater Fe Bed S301 15,472 2,732 bpd feed 0.18 $16,302,021 2014 i $16,302,021 0.75 4,440,161 1.5 4,440,161 4,440,161 $2,940,504 1
R-310 1 0 Hydrotreater Reactor, vessels, columns S301 15,472 2,732 bpd feed 0.18 $77,087,500 2014 i $77,087,500 0.75 20,996,224 1.5 20,996,224 20,996,224 $13,904,784 1K-311 1 0 Hydrogen Compressor S304-H2 3,786.0 4,926.0 lb/hr H2 1.30 $1,385,600 1Q 2011 b $1,385,600 0.8 1,710,365 1.1 1,881,402 1,895,882 $1,723,529 2
V-300 1 0 PSA for Hydrogen Recycle S425-H2 10 9.5 mmscfd H2 0.95 $1,750,000 2004 b $1,750,000 0.8 1,686,237 2.5 4,165,005 5,401,754 $2,186,945 3
A310 Total Subtotal
A310
HDO/N/S $33,576,724 $21,096,937
A350 Hydrocracking and Separations
R-350 1 0 Hydrocracker Unit + auxiliaries S338 15,472 1,021 bpd feed 0.07 $50,858,783 2014 i $50,858,783 0.75 6,619,706 1.5 6,619,706 6,619,706 $4,383,911 1
A350 Total Subtotal A350 HCK $6,619,706 $4,383,911
A400 Hydrogen Plant - OSBL
1 0
Stm Reformer system w/ associated
OSBL S425-H2 24.5 9.5 mmscfd H2 0.39 $50,198,438 2014 i $50,198,438 0.65 27,204,148 1.9 27,204,148 27,204,148 $14,168,827 4
A400
Total Subtotal
A400 H2
Plant $27,204,148 $14,168,827
A600 Power Generation - OSBL Costs accounted for in Hydrogen Plant except for turbine
E-601 1 0 Steam turbine + generator work 1,652 1,812 kW 1.10 $1,268,300 1Q 2014 b $1,268,300 0.85 1,371,980 1.2 1,632,657 1,641,203 $1,379,162 8
A600
Total Subtotal
A600 STM
Gen $1,641,203 $1,379,162
A700 OSBL - including cooling water system1 0 Wastewater treatment (NH3 stripping) S320+SCRB- 34,000 21,135 lb/hr 0.62 $1,217,800 1Q 2014 b $1,217,800 0.65 894,059 2.4 2,145,741 2,156,974 $898,739 81 0 Cooling Tower System S715 7,506,000 1,036,143 lb/hr 0.14 $260,852 2010 b $260,852 0.78 55,668 2.5 137,499 143,815 $58,225 61 0 Cooling Water Pump S715 7,001,377 1,036,143 lb/hr 0.15 $239,375 2007 b $239,375 0.3 134,943 2.1 288,778 316,644 $147,964 61 0 Makeup water pump S618+S710 80,411 33,964 lb/hr 0.42 $6,528 2007 b $6,528 0.3 5,041 4.7 23,792 26,088 $5,527 61 0 Plant Air Compressor S101(dry) X 10 262,454 61,113 lb/hr 0.23 $87,922 2007 b $87,922 0.3 56,784 1.6 89,150 97,753 $62,263 61 0 Firewater Pump S101(dry) X 10 262,454 61,113 lb/hr 0.23 $23,043 2007 b $23,043 0.3 14,882 3.7 55,064 60,377 $16,318 6
1 0 Instrument Air Dryer S101(dry) X 10 262,454 61,113 lb/hr 0.23 $8,349 2002 b $8,349 0.6 3,482 2.5 8,602 12,526 $5,071 6
1 0 Plant Air Receiver S101(dry) X 10 262,454 61,113 lb/hr 0.23 $21,005 2007 b $21,005 0.65 8,146 5.4 44,312 48,588 $8,932 6
1 0 Firewater Storage Tank S101(dry) X 10 262,454 61,113 lb/hr 0.23 $229,900 2007 b $229,900 0.65 89,154 1.5 130,165 142,726 $97,758 6
1 0 Feed Storage - 3 day S301 1,000,000 344,186 gallons 0.34 2,371,900 2010 b $2,371,900 0.7 1,124,221 1.4 1,573,909 1,582,148 $1,130,106 5
1 0 Product Storage - 3 day S392 1,000,000 83,952 gallons 0.08 2,371,900 2010 b $2,371,900 0.7 418,715 1.4 586,200 613,126 $437,947 5
1 0 Product Storage - 3 day S393 1,000,000 250,848 gallons 0.25 2,371,900 2010 b $2,371,900 0.7 900,915 1.4 1,261,281 1,319,216 $942,297 5
A700 Total Subtotal A700 OSBL $6,519,982 $3,811,148
Total Equipment Cost $75,561,763 $44,839,985
References
1 IHS PEP Yearbook 2014a - "Diesel from High Pressure Hydrocracking" (P=2000 psia)
2 Aspen Capital Cost Estimator, 7.3.2
3 Vendor Budget Estimate
4 IHS PEP Yearbook 2014b - "Hydrogen from Natural Gas by Steam Reforming"
5 Aspen Capital Cost Estimator, 7.3.1
6 Dutta et al. 2011
7 Kaiser and Gary 2007
8 Aspen Capital Cost Estimator, 8.8
C.1
- Compound Selection
C.1 HTL Liquid Products Composition
HTL organic products are a complex mixture of hundreds of compounds. The number and type of
compounds used in the Aspen model to represent HTL oil and the associated aqueous phase must
reasonably match key properties, such as CHONS, density, heating value, GC/MS data, expected HTL oil
distillation range, and aqueous solubility. The compounds chosen for the Aspen model are shown in
Table C.1. Note that this list does not imply that these compounds occur in the given percentages in
actual HTL oil, rather each compound represents a group of compounds that taken together exhibit the
bulk properties. Carbon dioxide and ammonia in the aqueous phase actually form their ionic species in
various amounts and types, including NH4+, NH2COO-, HCO3
-, CO32-. For simplification purposes, ion
formation is not simulated in the HTL model but their pure original compounds are considered.
Table C.1. Compounds used to model HTL liquid products.
HTL Oil
Heat & Mat'l
Balance Names Wt% C H O N S CAS
N-methylthiopyrrolidone C5H9NS 4.167% 5 9 1 1 10441-57-3
Toluene TOLUENE 2.986% 7 8 108-88-3
1H-Pyrrole, 3-ethyl-2,4,5-trimethyl- RYRO3ETM 5.284% 9 15 1 520-69-4
Phenol, 4-methyl- PHENO4M 0.459% 7 8 1 106-44-5
3-methyl-4-aminophenol AMIPHENO 0.919% 7 9 1 1 2835-99-6
Indole INDOLE 5.513% 8 7 1 120-72-9
3,7,11,15-tetramethyl-2-Hexadecene 2-PYTENE 6.892% 20 40 3452-07-1
1-Pentadecene C15OLEF 6.892% 15 30 13360-61-7
N-Methyldodecanamide MC12AMID 5.973% 13 27 1 1 27563-67-3
Palmitamide(Hexadecanamide) C16AMIDE 5.513% 16 33 1 1 629-54-9
9-Octadecenamide, (Z)- C18AMIDE 11.486% 18 35 1 1 301-02-0
Palmitic acid (n-Hexadecanoic acid) C16:0FA 7.811% 16 32 2 57-10-3
Oleic-acid C18:1FA 11.257% 18 34 2 112-80-1
Naphthalene, 1,2,3,4-tetrahydro-
1,1,6-trimethyl- C13H18 9.189% 13 18 91-20-3
Triphenylformazan HEVOIL1 9.915% 19 16 4 531-52-2
Dibenzyl-sebacate HEVOIL2 5.743% 24 30 4 140-24-9
100.000%
HTL Aqueous Phase wt% C H O N S CAS
Methanol METHANOL 2.347% 1 4 1 67-56-1
Ethanol ETHANOL 0.391% 2 6 1 64-17-5
Acetone ACETONE 0.196% 3 6 1 67-64-1
Acetic acid ACEACID 2.347% 2 4 2 64-19-7
propanoic acid PROACID 1.173% 3 6 2 79-09-4
CO2 CO2 68.178% 1 2 124-38-9
NH3 NH3 15.646% 3 1 7664-41-7
Ethyl-amine ETHAMIN 1.956% 2 7 1 75-04-7
Triphenylformazan HEVOIL1 2.738% 19 16 4 531-52-2
N-METHYLTHIOPYRROLIDONE 5.028% 5 9 1 1 10441-57-3
100.000%
C.2
C.2 Hydrotreated Oil Model Compounds
Similar to HTL oil, hydrotreated oil contains numerous compounds and a limited number are used for
modeling purposes. Table C.2 shows the mixture of compounds used to represent the hydrotreated oil in
the models. Note, that this list does not imply that these compounds occur in the given percentages in
actual hydrotreated oil, rather each compound represents a group of compounds that taken together
exhibit the bulk properties.
Table C.2. Compounds used to model hydrotreated product.
Compound
Heat & Mat'l
Balance Names C H O N Wt% CAS
Pentane N-PENTAN 5 12 4.47% 109-66-0
Pentane, 2-methyl- 2MPENTA 6 14 1.45% 107-83-5
Hexane HEXANE 6 14 1.12% 110-54-3
Cyclopentane, methyl- CC5-METH 6 12 0.78% 96-37-7
Hexane, 2-methyl- 2MHEXAN 7 16 1.45% 591-76-4
Cyclohexane, methyl- CC6-METH 7 14 0.56% 108-87-2
Piperidine PIPERDIN 5 11 1 0.22% 110-89-4
Toluene TOLUENE 7 8 1.68% 108-88-3
Heptane, 3-methyl- 3MHEPTA 8 18 1.12% 589-81-1
Octane OCTANE 8 18 0.78% 111-65-9
Cyclohexane, ethyl- ETHCYC6 8 16 0.84% 1678-91-7
Ethylbenzene ETHBENZ 8 10 0.84% 100-41-4
o-Xylene O-XYLENE 8 10 1.34% 95-47-6
Nonane C9H20 9 20 1.90% 111-84-2
Cyclohexane, propyl- CC6-PRO 9 18 1.45% 1678-92-8
Benzene, propyl- C3BENZ 9 12 0.89% 103-65-1
Nonane, 4-methyl- 4MNONAN 10 22 0.89% 17301-94-9
Decane C10H22 10 22 1.12% 124-18-5
Benzene, butyl- C4BENZ 10 14 1.68% 104-51-8
Benzene, 1-butenyl- C10H12 10 12 1.90% 824-90-8
Dodecane C12H26 12 26 1.34% 112-40-3
Naphthalene, 1,2,3,4-tetrahydro- 1234NA 10 12 2.24% 119-64-2
Cyclohexane, hexyl- CC6-HEX 12 24 2.24% 4292-75-5
Naphthalene, 1,2,3,4-tetrahydro-
1,4-dimethyl- 14DMNAPH 12 16 2.80% 4175-54-6
Benzene, heptyl- C7BENZ 13 20 3.36% 1078-71-3
Benzene, octyl- C8BENZ 14 22 2.57% 2189-60-8
Pentadecane C15H32 15 32 2.24% 629-62-9
Hexadecane C16H34 16 34 6.71% 544-76-3
Heptadecane C17H36 17 36 11.19% 629-78-7
2,6,10,14-Tetramethyl-
hexadecane TETMC16 20 42 6.71% 638-36-8
Octadecane C18H38 18 38 10.07% 593-45-3
Eicosane C20H42 20 42 6.71% 112-95-8
Palmitic acid C16:0FA 16 32 2 1.57% 57-10-3
Heneicosane C21H44 22 46 1.68% 629-94-7
Tetracosane C24H50 24 50 3.91% 646-31-1
Di-2-ethylhexyl-adipate C22H42O4 22 42 4 2.24% 103-23-1
Octacosane C28H58 28 58 2.80% 630-02-4
Triacontane C30H62 30 62 0.89% 638-68-6
Dotriacontane C32H66 32 66 2.24% 544-85-3
D.1
- Indices
Table D.1. Labor indices.
Source: Bureau of Labor Statistics
Series ID: CEU3232500008 Chemicals
Average Hourly Earnings of Production Workers
Current indices @ http://data.bls.gov/cgi-bin/srgate
YEAR INDEX YEAR INDEX
2005 19.67 2010 21.07
2006 19.60 2011 21.46
2007 19.55 2012 21.45
2008 19.50 2013 21.40
2009 20.30 2014 21.49
Table D.2. Capital cost indices.
Chemical Engineering Magazine, CEI annual index YEAR INDEX YEAR INDEX
1990 357.6 2004 444.2
1991 361.3 2005 468.2
1992 358.2 2006 499.6
1993 359.2 2007 525.4
1994 368.1 2008 575.4
1995 381.1 2009 521.9
1996 381.7 2010 550.8
1997 386.5 2011 585.7
1998 389.5 2012 584.6
1999 390.6 2013 567.3
2000 394.1 2014 576.1
2001 394.3 2015 556.8
2002 395.6 2016 541.7
2003 402.0
D.2
Table D.3. Inorganic chemical indices.
Source: SRI International Chemical Economics Handbook, Economic Environment of the Chemical
Industry, September 2006
Current indices @https://www.sriconsulting.com/CEH/Private/EECI/EECI.pdf
YEAR INDEX YEAR INDEX
1990 123.6 2003 164.6
1991 125.6 2004 172.8
1992 125.9 2005 187.3
1993 128.2 2006 196.8
1994 132.1 2007 203.3
1995 139.5 2008 228.2
1996 142.1 2009 224.7
1997 147.1 2010 233.7
1998 148.7 2011 252.1
1999 149.7 2012 260.3
2000 156.7 2013 263.9
2001 158.4 2014 269.2
2002 157.3 2015 264.8
E.1
– TCLP Testing Results
Table E.1. TCLP analysis for HTL solids from Detroit sludge (methods 6010B for metals and 7470A for
mercury).
HTL Solids TCLP Analysis
Metal WW-06 BC Comp, mg/L Federal Limits, mg/L
As n.d. 5
Ba 0.11 100
Cd n.d. 1
Cr n.d. 5
Pb 0.12 5
Hg n.d. 0.2
Ni 0.48 n/a
Se n.d. 1
Ag n.d. 5
Zn 16.7 n/a
F.1
- Process Efficiency Metrics
In addition to setting technical and economic targets, BETO determines process efficiency metrics
that can be used to estimate environmental sustainability of biofuels/products pathways, including product
yield, carbon efficiency, energy inputs, and water consumption. These metrics are used in supply chain
sustainability analyses for the pathways using the GREET (ANL 2016a) and WATER (ANL 2016b)
models. Table F.1 lists metric values for the HTL plant (sludge to biocrude) and the overall sludge to
upgraded fuel conversion process. Solids reduction efficiency is also listed for the sludge HTL process as
it is of particular interest to WWTPs for any sludge treatment strategy under consideration.
Table F.1. Process efficiency metrics for sludge HTL and biocrude upgrading goal case.
Process Efficiency Metric
Sludge to
Biocrude
Overall Sludge to
Fuel Blendstocks
Product Yield (% w/w of feedstock)a 48% 39%
Carbon Efficiency of Conversion Process
% C in fuel / feedstock
% C in fuel / feedstock + natural gas
72%
66%
64%
53%
Net Electricity Import (kWh/gge product) 1.6 2.0
Purchased Natural Gas Import (million Btu/gge product)b 0.02 0.04
Product Energy / Input Energy (LHV)c 4.2 2.8
Water Consumption (gal/gge product) 0.05 1.2
Solids Reduction, %
Dry basis
Dry, ash-free basis
76%
89%
n/a
(a) On a dry, ash-free basis
(b) Calculated using natural gas lower heating value of 21,152 Btu/lb from process model.
(c) Includes only direct inputs of electricity and natural gas (for HTL and THROX unit). Biocrude
transport energy not included.
References
ANL. 2016a. Greenhouse Gases, Regulated Emissions, and Energy Use in Transportation (GREET)
Model. Argonne National Laboratory, Argonne, IL. .
ANL. 2016b. Water Analysis Tool for Energy Resources (WATER). Argonne National Laboratory,
Argonne, IL.
